PB 198 809
APPLICABILITY OF CATALYTIC OXIDATION TO THE
DEVELOPMENT OF NEW PROCESSES FOR REMOVING
S02 FROM FLUE GASES - VOLUME II --EXPERIMENTA L
PROGRAM
R. E. Opferkuch, et al
January 1971
NATIONAL TECHNICAL INFORMATION SERVICE
Distributed .,. 'to foster, serve
and promote the nation's
economic development
and technological
advancement.'
U.S. DEPARTMENT OF COMMERCE
This document has been approved lor public release and sale.
-------
APPLICABILITY OF CATALYTIC OXIDATION
TO THE DEVELOPMENT OF NEW PROCESSES
FOR REMOVING S02 FROM FLUE GASES
Volumell- EXPERIMENTAL PROGRAM
Contract No. PH 22-68-12
Prepared by
R.E. Opferkuch
S.M. Mehta
M.G. Konicek
D.L. Zanders
Submitted to
. Division of Process Control Engineering
National Air Pollution Control Administration
Environmental Health Services
U.S. Public Health Service
U.S. Department of Health. Education, and Welfare
5710 Wooster Pike
Cincinnati, Ohio 45277
-------
BIBLIOGRAPHIC DATA
SHEET
). Report No.
APTD-0676
3. Recipient's Accession No.
5. Report Date
January 1971
4. Title and Subtitle
Applicability
of Catalytic Oxidation to the
ment of New Processes for Removing SO From
Gases Volume II - Experimental Program
Develop
Flue
6.
7. Author(a)IT E. Opferkuch (Project Leader)
S. M. Mehta, M. G. Konicek, D. L. Zanders
8. Performing Organization Kept.
No.
>. Performing Organization Name and Address
Monsanto Research Corporation
Dayton Laboratory
Dayton, Ohio 45407
10. Project/Task/Work Unit N-
11. Contract/Grant No.
PH 22-68-12
11 Sponsoring Organization Name and Address
Process Control Engineering Program
National Air Pollution Control Administration
Environmental Health Service, U.S. Dept. of HEW
5710 Wooster Pike
Cincinnati, Ohio 45277
13. Type of Report 4t Period
Covered
14.
IS. Supplementary Notes
16. Abstracts
•During the first phase of this project, which covered the identification
and evaluation of existing and potential methods of applying catalysis
to the oxidation and removal of S0_ from power plant stack gas, a large
quantity of information and data was accumulated and assessed. This
evaluation revealed the need for laboratory verification of the publishe|d
data on promising oxidation systems. This volume presents subsequent
data and information generated in the laboratory and on the drawing
board.
17. Key Words and Document Analysis. 17o. Descriptors
Tests
Catalysis
Oxidation
Catalytic Converters
Sulfur dioxide
Flue gases
Electric power plants
I7b, Identifiers/Open-Ended Terms
17e. COSATI Field/Group
13/B
IB. Availability Statement
""limited
19.. Security Class (This
Repon)
UNCLASSIFIED
20. Security Class (This
UNCLASSIFIED
21. No. of Pages
115
22. Price
FOUM NTIC-tB 110-70)
USCOMM-DC 40820-P7I V!
-------
DISCLAIMER
This report was furnished to the Air Pollution
Control Office by Monsanto Research Corporation
Dayton Laboratory
Dayton, Ohio 45407
in fulfillment of Contract No. PH 22-68-12
-------
MRC-DA-245
APPLICABILITY OF CATALYTIC OXIDATION TO
THE DEVELOPMENT OF NEW PROCESSES FOR
REMOVING S02 FROM FLUE GASES
Volume II - Experimental Program
Contract No. PH 22-68-12
MRC Job No. 6708
Prepared by
R. E. Opferkuch, Project Leader
S. M. Mehta
M. G. Konicek
D. L. Zanders
MONSANTO RESEARCH CORPORATION
DAYTON LABORATORY
Dayton, Ohio
January 1971
Submitted to
Process Control Engineering Program
National Air Pollution Control Administration
Environmental Health Service
U.S. Department of Health, Education, and Welfare
5710 Wooster Pike
Cincinnati, Ohio 45277
-------
FOREWORD
The Intent of this volume is to present in an organized manner
the accumulation and assessment of all experimental and design work
accomplished during the execution of Contract PH 22-68-12 by
Monsanto Research Corporation, Dayton, Ohio. In that this volume
represents the second phase of a three-phase program of concurrent
tenure, occasional reference to phases I and III may be seen at
various points through the text. Further definition of these areas
may be found in the respective volumes, each of which is reported
under separate cover.
This final report is presented in three volumes in an effort to
make the material accessible on the assumption that it is of
practical value and therefore will be put to use.
Volume I is intended to contain all, and only, that material derived
from, or related to, the literature search. Essentially all
information in Volume I is directly based on the literature.
Volume II presents data and information generated in the laboratory
and on the drawing board.
Volume III is an Indexed bibliography.
Finally, guidance through the three volumes is offered in the form
of the Foreword, Project General Summary and Tables of Contents in
each of the three volumes.
The authors wish to acknowledge the many helpful comments and
suggestions of the NAPCA Project Officer, Mr. George L. Huffman.
ii
• MONSANTO RESEARCH CORPORATION •
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PROJECT GENERAL SUMMARY
The salient features of the project are summarized briefly below.
Expansion and details are given in the texts of Volumes I and II.
The main objectives of the program were:
a. Search all available literature for pertinent information
relative to the catalytic oxidation of sulfur dioxide, i.e.,
active materials, mechanisms of catalysis, methods of
application, equipment employed, etc. Identify, describe
and evaluate processes disclosed in the literature to have
commercial potential for removal of sulfur dioxide from
flue gas by oxidation.
b. Test, in the laboratory, candidate materials and methods
suggested in the literature for potential application to
removal of sulfur dioxide from flue gas by catalytic
oxidation.
c. Identify at least one effective catalyst for the desired
application and design a process for removal of sulfur
dioxide from flue gas by catalytic oxidation and recovery
of the sulfur value.
An intensive search of the literature revealed the following:
a. The transition metal oxides, notably vanadia, and platinum
were the most commonly employed solid catalysts for practical
conversion of sulfur dioxide to trioxide. Nitrogen dioxide
was the only practical gaseous catalyst noted.
b. Kinetic equations describing conversions of sulfur dioxide
over vanadia or platinum catalysts were derived from data
relative to commercial production of sulfuric acid, i.e.,
high concentrations of sulfur dioxide. There was nothing
available to describe results at the comparatively low
concentrations of sulfur dioxide found in flue gas.
c. A number of processes were described as having commercial
potential for flue gas cleaning. Comparative cost-
performance evaluation of these oxidation processes
eliminated all but two types as having realistic commercial
potential, viz., one type based on vanadia catalyst and one
based on nitrogen dioxide catalyst.
d. The most practical mode of recovery of oxidized sulfur value
from a use standpoint, in this country, is production of
fertilizer grade sulfuric acid.
• MONSANTO RESEARCH CORPORATION •
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Laboratory tests and comparative,evaluation of commercial and
experimental catalysts indicated the following:
a. Commercial vanadia catalysts employed in production of
sulfuric acid are effective in converting sulfur dioxide
at concentrations found in flue gas.
b. Although platinum catalyst performance is essentially
equivalent to that of vanadia, a cost-performance
comparison indicated vanadia is ten times more effective.
c. No other candidate materials were shown to be as
effective as vanadia, platinum or nitrogen dioxide in
converting sulfur dioxide to trioxide,
d. Nitrogen dioxide was the only practical "low temperature"
catalyst observed.
e. Using nitrogen dioxide a,s a catalyst, It is potentially
practical to remove botjh sulfur dioxide and indigenous
nitrogen oxides from flue gas simultaneously.
From preliminary process designs, and cost estimates, based on
laboratory data generated in this program, the following emerge:
a. Processes, based on vanadia catalyst, for oxidizing
sulfur dioxide and removing it from power plant flue gas,
as sulfuric acid, are likely to cost in the range of $12
to $25 of capital per Installed Kw of power plant capacity.
A large portion of the cost results from the need for
corrosion resistant equipment.
b. Operating costs for vana,dia based propesses are likely
to be in the range of P,5Q to 0.7*1 mills/Kw-Hr generated
before sulfur value net back.
c. A substantial reduction in capital and operating costs
are potentially available through a technique of sorbing
the oxidized product gas from the main flue gas stream
and recovering it separately.
d. The vanadia based processes are better suited to proposed
new power plant installations than to existing plants
because of numerous difficulties in retro-fit to existing
power plants.
iv
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VOLUME II
TABLE OF CONTENTS
Page
INTRODUCTION 1
I. EXPERIMENTAL EQUIPMENT 3
A. Catalyst Test Unit 3
B. Temperature Control 3
C. Reactor 7
D. Analysis System 7
II. CATALYST PREPARATION 19
III. CATALYST TESTING PROCEDURE 23
IV. CATALYST TEST RESULTS AND DISCUSSION 27
A. Commercial Vanadia Catalysts 27
B. Commercial Platinum Catalyst 33
C. Experimental Catalysts 33
D. Molecular Sieve Adsorption Studies 41
E. S02 Removal Characteristics of "Red Mud" 42
V. DEVELOPMENT OF A SIMULTANEOUS SOX-NOX REMOVAL PROCESS 45
A. Initial Laboratory Jnvestigations • 45
B. Vapor Pressure Studies on the System 48
C. Absorption Studies of NO and N02 into Sulfuric
Acid • 50
D. Proposed SOX-NOX (SONOX) Removal Process
Description 54
1. General Description 54
2. Trifunctional Absorption Tower . 56
a. Bottom Section of Absorption Tower 56
b. Middle Section of Absorption Tower 56
c. Top Section of Absorption Tower 57
3. Acid Demister 57
4. Stripper 57
5. NO Oxidizer 58
6. Acid-Acid Heat Exchanger 58
7. Heat and Material Balances 58
E. Comparison Between Tyco and Monsanto Modified
Chamber Process 58
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TABLE OF CONTENTS (Cont'd)
1. General Description of Process Plow
Differences 58
2. Detailed Process Comparison 60
3. Detailed Comparison Using Identical Strippers 62
F. Remaining Questions to be Answered 62
VI. PROCESS DESIGN ' 65
A. General 65
B. Case I. MRC/NAPCA Process - New 1400 Mw Plant 69
C. Case I-A. Reversible Dry Absorption of SQ3 83
D. Case II - MRC/NAPCA Process - Existing 220 Mw
Plant 92
E. Case III - Small Copper Smelter 105
Appendix I - Commercial Catalyst Test Data 115
Appendix II - Experimental Catalyst Test Data 129
Appendix III - Effect of Catalyst Particle Geometry on
Pressure Drop in a Fixed Bed > 135
Appendix IV - Reactor Design . .
Appendix V - Relationship of W/F Factors and
Space Velocity
Appendix VI - Comparative Cost of Platinum and
Vanadia Catalysts 153
Appendix VII - Sorptlon Isotherms for Molecular Sieves l6l
Appendix VIII - Tyco Modified Chamber Process Report 171
Appendix IX - Vapor Pressure Apparatus and Procedures 183
Appendix X - S03-Product Acid Relationships 187
vi
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LIST OF FIGURES
Pag(
1 Test Unit Schematic Plow System 4
2 Catalyst Test Unit Control Panel 5
3 Plow Controls 6
4 Reactor Furnace Controls 8
5 Glass Reactor and Preheater Section 9
6 Catalyst Bed and Temperature Sensing Assembly • 10
7 Reactor Furnace Enclosure 11
8 Reaction Section Components 12
9 Furnace Sub-assembly 13
10 Reactor and Furnace Enclosure 14
11 Process Stream Cooler Between Reactor and Chromatograph 16
12 Process Stream Analyzer 17
13 Analysis System Programmer 18
14 Catalyst Pelletizer 20
15 Catalyst Pellet Surface (6000X) Before and After
Heating at 800°F 21
16 Comparison of Experimental and Predicted Effects of
W/F on Conversion at 800°F 26
17 W/F Conversion Profiles for Catalyst 'A' 28
18 W/F Conversion Profiles for Catalyst-E . 29
19 Comparative Conversion Efficiency of Two Commercial
Vanadia Catalysts at 800°F 30
20 Comparative Conversion Efficiency of Two Commercial
Vanadia Catalysts at 850°F 31
21 Comparative Conversion Efficiency of Two Commercial
Vanadia Catalysts at 900°F 32
22 W/F Conversion Profiles for Commercial Platinum Catalyst 34
23 Comparative Conversion Efficiency of Platinum and
Vanadia Catalysts at 900°F 35
24 Effect of Time on Conversion Efficiency of Catalyst-B
at 800°F, Constant Reactor Conditions (Run No. 6) 36
25 Effect of Time on Conversion Efficiency of Catalyst-B
at 800°F, Constant Reactor Conditions (Run No. 7) 37
26 Conversion Efficiency versus Time Profile @ Constant
Operating Conditions 43
vii
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LIST OF FIGURES (Cont'd)
Page
27 Laboratory Arrangement for NOX Scrubber 46
28 Experimental Apparatus for Vapor Pressure Studies 49
29 Adjusted Vapor Pressure of NO and N02 Over Nitrose 51
30 Absorption Studies Apparatus 52
31 Process Design Schematic 55
32 W/F Conversion Profiles for Catalyst 'A' 68
33 MRC/NAPCA Process, Large New Power Plant Facility,
1400 Mw, Process Plow Diagram 70
34 MRC/NAPCA Process, Large New Power Plant Facility,
1400 Mw, Plot Plan 71
35 MRC/NAPCA Process, Large New Power Plant Facility,
1400 Mw, Elevation Drawing 72
36 MRC/NAPCA Process, Large New Power Plant Facility,
1400 Mw, Schematic Plan 76
37 MRC/NAPCA Process, Large New Power Plant Facility,
1400 Mw, Instrument Flowsheet 77
38 Effect of Product Credit on Operating Cost of MRC/-
NAPCA Process (Large New Power Plant Facility) 84
39 Reversible Dry Absorbent Process, Process Flow Diagram 90
40 Effect of Product Credit on Operating Cost of
Reversible Dry Absorbent Process 97
4l MRC/NAPCA Process, Small Existing Power Plant Facility,
220 Mw, Process Flow Diagram . 98
42 Effect of Product Credit on Operating Cost of MRC/-
NAPCA Process (Small Existing Power Plant Facility) 104
43 MRC/NAPCA Process, Smelter Facility 107
44 Effect of Product Credit on Operating Cost of MRC/-
NAPCA Process (Smelter Facility) 113
45 W/F-Conversion Profile for Catalyst 'A' at 900°F H6
46 W/F-Conversion Profile for Catalyst 'A1 at 850°F 117
47 W/F-Converslon Profile for Catalyst 'A' at 800°F 118
48 Effect of W/F on Conversion with Catalyst-B at 900°F 120
49 Effect of W/F on Conversion with Catalyst-B at 850°F 121
50 Effect of W/F on Conversion with Catalyst-B at 800°F 122
51 Effect of W/F on Conversion with Catalyst-E at 900°F 124
viii
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LIST OF FIGURES (Cont'd)
Page
52 Effect of W/F on Conversion with Catalyst-E at 850°F 125
53 Effect of W/F on Conversion with Catalyst-E at 800°F 126
54 Effect of Catalyst Particle Geometry on Pressure Drop
in a Fixed Bed 139
55 Sorption Isotherms for Molecular Sieve, SK-20 163
56 Sorption Isotherms for Molecular Sieve, SK-400 16^
57 Sorption Isotherms for Molecular Sieve, SK-400 165
58 Sorption Isotherms for Molecular Sieve, SK-410 166
59 Sorption Isotherms for Molecular Sieve, SK-410 167
60 Sorption Isotherms for Molecular Sieve, 13X 168
61 Sorption Isotherms for Molecular Sieve, SK-110 16.9
62 Baseline Process 175
63 Isothermal Scrubber 176
64 Catalytic Stripper 179
65 Catalytic Chamber Process 180
66 Vapor Pressure Apparatus and Procedures 184
67 Effect of S03 Concentration in Flue Gas on
Dew Point of the Acid 188
68 Relationship of Flue Gas Dew Point to Acid
Concentration in Condensate 189
ix
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LIST OF TABLES
Page
1 Catalyst Bulk Density 24
2 Catalyst Test Data 25
3 Summary of Economics for Vanadia and Platinum Catalysts 38
4 Effect of N02 on Total Conversion of Sulfur Dioxide 47
5 Summary of Laboratory N203 Absorption Studies 53
6 Heat and Material Balance - High Temperature Stripper 59
7 Comparison of Monsanto and Tyco Processes (As Designed) 6l
8 S02 Removal from Flue Gas 73
9 S02 Removal from Flue Gas 74
10 Catalytic•Converter 78
11 Acid Recovery and Mist Collection Tower 79
12 Acid Cooler 80
13 Acid Pump 8l
14 Induced Draft Fan 82
15 Capital Cost Estimate Summary 85
16 Equipment Cost Estimate Summary 86
17 Working Capital Estimate Summary 87
18 Operating Cost Estimate Summary 88
19 S02 Removal from Flue Gas 91
20 Capital Cost Estimate Summary 93
21 Equipment Cost Estimate Summary 94
22 Working Capital Estimate Summary 95
23 Operating Cost Estimate Summary 96
24 S02 Removal from Flue Gas 99
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LIST OF TABLES (Cont'd)
Page
25 Capital Cost Estimate Summary 100
26 Equipment Cost Estimate Summary 101
2? Working Capital Estimate Summary 102
28 Operating Cost Estimate Summary 103
29 S02 Removal from Flue Gas 108
30 Capital Cost Estimate Summary 109
31 Equipment Cost Estimate Summary 110
32 Working Capital Estimate Summary 111
33 Operating Cost Estimate Summary 112
34 Catalyst - A Test Results 119
35 Catalyst-B Test Results 123
36 Catalyst-E Test Results 127
37 Experimental Catalyst Test Data 130
38 Effect of Catalyst Particle Geometry on Pressure
Drop in a Fixed Bed 140
39 Optimum Diameter of a Reactor Bed 1^8
40 Summary of Economics for Vanadia and Platinum
Catalysts 155
xi
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VOLUME II SUMMARY
During the first phase of this contract, which covered the
Identification and evaluation of existing and potential methods of
applying catalysis to the oxidation and removal of S02 from power
plant stack gas, a large quantity of information and data was
accumulated and assessed. This evaluation revealed the need for
laboratory verification of the published data on promising oxidation
systems. While the literature noted a large number of Individual
materials, and combinations of materials, as being capable of
converting S02 to 803, the experimental conditions, conversion
efficiencies, and extrapolated economics cited for nearly all cases
were impractical with respect to the flue gas application.
Essentially all these studies were related to conversion with
vanadla-based catalyst in gas streams containing 5-12$ S02, typical
of that in contact with sulfuric acid plant operation. Accordingly,
it became necessary to provide laboratory verification of these data
under simulated flue gas conditions.
Commercially available catalysts were tested over a broad range of
conditions, resulting in the conclusion that vanadia-based materials
were the only practical compositions presently available with the
capability to remove SOa from power plant flue gas under normal
operating conditions.
A large number of potentially promising experimental compositions
were tested, but none was found which exhibited properties superior
to that of commercial vanadia-based catalysts for the pertinent task.
During this experimental investigation enough promising data was
gathered to merit further research in the area of new catalysts
which would operate at a lower temperature than those presently
available.
The development of a process which will simultaneously remove both
S02 and NOx was implemented experimentally as a logical outgrowth
of the experimental program. A detailed description of this process
has been included in the text (Section V). As this process bears
some similarity to the Tyco removal process, a comparison between
the two is also shown.
Preliminary catalytic oxidation process designs were developed for
application to a large new power plant facility, a large existing
power plant facility, a small existing power plant facility, and a
new smelter facility. These designs include estimates of investment
and operating costs. Since a great deal of data was obtained during
the execution of the experimental and design phases of the program,
most of it has been relegated to an appendix section at the back of
the volume.
xli
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VOLUME II CONCLUSIONS
1. Vanadia catalysts available for commercial production of sulfuric
acid are effective in converting SO?, in flue gas to SO3.
2. Platinum catalysts available for commercial hydrocarbon reforming
are effective in converting S02 in flue gas to SO3.
3. -Although platinum catalyst is essentially equivalent in
performance to vanadla catalyst, it is an order of magnitude
more expensive.
4. The presence of NOX in the flue gas stream had no apparent
effect on catalyst conversion efficiency.
5. No catalyst compositions tested, other than platinum and
vanadia based compositions, were deemed satisfactory because
of their instability to either SO,, or S03<
6. No other combination of catalyst-promoter tested was as effective
as the vanadia-based compositions for oxidizing S02 to S03.
7. Catalytic oxidation processes are better suited to new plant
design because of high costs associated with retro-fit into
existing power station designs.
8. Fleeting evidence in support of a proposed mechanism in the
literature for vanadia catalysis, was obtained experimentally.
9. No low temperature catalyst, other than nitrogen dioxide, was
found'during the1 experimental search - however, enough evidence
was gathered to merit further investigation into low temperature,
solid catalysts.
xiii
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INTRODUCTION
The theories and principles of catalysts are discussed in Volume I
of this report. The most significant conclusion emerging from the
discussions in Volume I was that heterogeneous catalysts by vandaia
is the most effective means of converting S02 to S03. This state-
ment, however, does not specify a formulation of vanadia mixed. with
anything else and does not suggest, in itself, how the vanadia is con-
tacted with the sulfur dioxide. Furthermore, there were, in the litera-
ture, several strong Indications, or leads, to potentially new cata-
lysts that might operate at lower temperatures, or' that might be more
active per unit, weight. . .
The experimental effort, then, had two major objectives: (1) to seek
"low" temperature or more active catalysts and (2) to define at least
one catalyst around which a process could be designed for removing
S02 from power plant stack gas.
In the following discussion, it is helpful to understand .the general
terminology and structure of solid catalysts. The basic parts of
a solid catalyst consist of:
1. Catalyst - active material, usually noble
metals or transition metal oxides.
2. Promoter - usually an inorganic salt or
salt mixture that enhances the activity
of the catalyst.
3. Support - an inert material such as Fuller's
earth or alumina used to provide bulk to the
mixture and increase surface area.
The relative proportions of the three materials are by no means
fixed but fall, roughly, into the following ranges, by weight:
Catalyst -
-------
The ingredients can be combined by:
1. Dry mixing, pelletizing, calcining.
2. Adding enough liquid to make a dough and extrud-
ing pellets, drying, calcining.
3. Dry mixing catalyst and promoter, calcining,
grinding the "clinker", dry mixing with support,
pelletizing.
4. Impregnating support with a solution of catalyst
and promoter, drying, pelletizing, calcining.
It may not always be required to calcine. When it is necessary,
however, the purpose is usually to convert the catalyst-promoter
mixture to the optimally active composition. For example, the
desired composition of the promoter in a vanadia catalyst is the
pyrosulfate, whereas the promoter may initially be added as the
sulfate.
The number of permutations and combinations of catalyst pretreat-
ment operations promotes the impression that the field of catalysis
is a "black art".
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I. EXPERIMENTAL EQUIPMENT
A. Catalyst Test Unit
The system that was designed and fabricated for the purpose of test-
ing catalysts consists of the following functional sections':
Preparation and metering of simulated flue gas
Temperature control
Reactor
Analysis
Figure 1 is a schematic flow diagram of .the test stand. Figure 2
is a view of the control panel.
The gas preparation section is designed to permit mixing and feeding
of simulated flue gas from cylinders, including calibrating gas mix-
tures. Precise flow control is necessary here, and is provided through
multistage pressure regulation, precise flowmeters, and very sensitive
flow controllers.
The sample gas flow-control subassemblies, as.shown in Figure 3» con-
sist of two-stage regulation at the cylinder source followed by
rotameters with needle valve control. Flow controllers are installed
downstream of the rotameters and are designed to control from con-
stant upstream pressure sources through variable.downstream-pressure
sources. The flow rate set point is manually set by adjusting an
external valve handle. This valve acts as a variable orifice to
provide a pressure drop across the high and low pressure chambers
of the diaphragm. If the pressure .drop changes due to a change
in supply or outlet condition, the diaphragm will change to the operat-
ing point of an internal throttling valve to re-establish the pressure
drop across the manual valve. The flow control system has the
capability of metering individual component flows, as well as multi-
component flows. In the case of individual component flow, the
components are mixed downstream of the flow metering assembly through
a specially designed mixing chamber.
B. Temperature Control
The mixed gas stream next passes through .a section of preheated line
prior to entrance into the reactor section, Tn the reactor section,
the gas is heated even further in an eight foot coil within the
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Figure 1. Test Unit Schematic Plow .System
o
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a
m
v>
m
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TJ
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L oo<5<5
066
A, Gas Bottles
B. Flow Control Valves
C. Flow Meters
D. Pressure Gauges
E. Flow Controllers
F. Mixer
G. Inlet Gas Sample
H. Reactor Bypass
I. Preheat Section
J. Furnace Preheat Section
K. Catalyst Bed
L. Process Stream Cooler
M. Condensables Trap
N. Process Stream Vent
0. Process Stream Analyzer
P. Exhaust Stack
-------
Figure 2. View of Test Unit Control Panel
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•
i
.
„
« . •
•
I.
'
-
.
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.
•
.
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.
.
.
:
Figure 3. Plow Controls
• MONSANTO RESEARCH CORPORATION
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reactor furnace enclosure. Thermal input is controlled by two 220-
volt, 3-phase Variacs, one 3-phase contactor and a West temperature
controller, allowing exact adjustment of input power and temperature
control within +1°F. Figure 4 shows the temperature control assembly.
C. Reactor
Gas from the flow control system enters the reactor, Figure 5, through
an eight-foot spirally-wound,heating section ahead of the reaction tube
The hot gas then flows through the catalyst bed and out the exit•arm.
Figure 6 shows an enlarged schematic of the catalyst bed, thermo-
couples and supporting assembly,all of which may be positioned in,
or removed from, the reactor as one unit, allowing practical inter-
change of catalyst beds. Figure 7 details the reactor-assembly
furnace-enclosure. The enclosure is mounted on a 3/4-inch, 100-lb
steel plate which serves as a basic mounting surface as well as pro-
viding stability through a low center of gravity. A high-temperature
blower extends through the base plate and into a stainless steel tank
12 inches in diameter and 24 inches in length. Within this tank,
twelve 500-watt, high-temperature, strip heaters form an annulus sur-
rounding the glass reactor tube and provide .an 18-inch heated zone
along the length of the reactor and preheat assemblies. Enclosed
in a squirrel cage housing, the fan blade exhausts up the outside
of the heat bank, and intakes along the center axis, providing a
uniform thermal environment for the catalyst bed. The primary
tank enclosure is surrounded by a minimum of three inches of
FIBERFRAX ceramic insulation contained by a 30-gallon steel drum
mounted on the base plate. Components of the reactor section are
shown in Figures 8, 9 and 10.
D0 Analysis System
The analytical system consists of an automatic, chromatographic pro-
cess stream analyzer. The analyzing system has the capability to
resolve accurately S02, 02, and C02, with a multistage range
selector for S02, to encompass 0-0,2%, 0-0.4%, and Q-5% of this
component. Other ranges are easily incorporated with the addition
of modular, attenuator cards in the system programmer0
The complete analyzing system consists of four major units: sample
conditioner, analyzer, programmer, and read-out device consisting
of an L&M series H recorder,,
The sample conditioner receives the raw sample from the process
stream and prepares it for introduction into the analyzer. The
chromatographic analyzer separates the sample into its individual
constituents and, as they elute, transmits signals to the pro-
grammer. Signals are proportional to the concentrations of the
various components. The analyzer utilizes a thermal .conductivity
detector with four filaments.
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Figure 4. Reactor Furnace Controls
8
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BASKET a T.C. ASSEMBLY
\
SAMPLE GAS
OUT
PREHEATER
SECTION
SAMPLE GAS
IN
CATALYST
BED
figure 5. Glass Reactor and Preheater Section
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REACTOR TUBE
CAP
EXIT GAS
TEMPERATURE
CATALYST
BASKET
(25 mm O.D. x 40 mm hgt)
SCREEN '
INLET GAS
TEMPERATURE
BASKET
UPPORTING
ROD
CATALYST
BED
(7/8 In. dia. x 1.5 in. hgt)
Figure 6. Catalyst Bed and Temperature Sensing Assembly
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INSULATION
CONTMNER
HEATER
BANK
REACTOR
ENTRANCE COLLAR
TANK
SUPPORTS
HEAT
MJMDf
HEATER WIRE
FEED THRU
CONTROLLED
ENVIRONMENT
CHAMBER
MOUNTING
PLATE
HIGH TEMPERATURE
FURNACE MOTOR
Figure 7- Reactor Furnace Enclosure
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•
*
*•
I
fl' ""•;
i
1
*r'
'
•
•
•
.
I
•
'
*~
'•-!''•£& t
Figure 8. Reaction Section Components
12
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•
.
•
.
•
•
f
Figure 9. Furnace Sub-assembly
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Figure 10. Reactor and Furnace Enclosure
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The programmer provides the control signals required to operate
the analyzer, and converts the analyzer-output signal into a
form compatible with the read-out device. The sampling pro-
grams offered are:
• Reactor inlet/outlet, sequential, iterative
• Iterative inlet only
* Iterative outlet only
* Manual select
Various components of the analytical system are shown in Figure.s
11 through 13.
15
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Figure 11. Process Stream Cooler
Between Reactor and Chromatograph
x
/*?
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Figure 12. Process Stream Analyzer
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Figure 13. Analysis System Programmer
KEPRODUCIBU
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II. CATALYST PREPARATION
Laboratory equipment obtained for catalyst preparation Included
a large capacity oven (0-300°F) for drying samples, a small
high-temperature furnace (0-l830°F) for calcining operations;
and a pelletizing press (Manesty), Figure 14.
The primary method of sample preparation employed in catalyst
studies consisted of the following:
• Weigh out constituents in fine powder form
(pregrind if necessary)
• Homogeneously dry mix constituents
* Bake mixture overnight to remove free water
' Pelletize
• Calcine
The support material most commonly used was calcium montmorillonite,
or Fuller's earth. It was checked for its capacity to retain
molten alkali salts (promoters) by preparing a blend of it with,
for example, potassium blsulfate (20% by weight), crushed (to
approximately 300 mesh. This blend was then pelletized arid heated
in a furnace 120°F above the normal melting point of the KHSOi,.
Subsequently, the pellets were removed and examined under magnifi-
cation, for signs of running or sweating of the KHS04.
In Figure 15, obtained with .the aid of the scanning electron micro-
scope, the upper photograph (6000X) shows the surface of a catalyst
pellet before calcining. The large particles, roughly 3 microns
in size, are, we believe, the mixture of catalyst and promoters.
The smaller particles can clearly be identified as diatomaceous
earth, the support. The surface exposed to gaseous reactants is
highly Irregular and the pores exist as the space between or
within the particles.
The lower photograph (6000X) in Figure 15 shows the surface of a
pellet after calcining at 800°F. The large particles are absent
and the support material appears to have been rather uniformly
coated, presumably by the molten promoters in which the catalyst
is dispersed.
In some preparations, alumina was employed as support in a manner
similar to that for Fuller's earth.
19
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Figure 14. Catalyst Pelletizer
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Figure 15.
Catalyst Pellet Surface
(6000X) Before (Upper)
and After (Lower) Heating
at 800°F
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A second modification of preparation consisted of mixing catalyst
and promoter, calcining, grinding the "clinker" and dry mixing
with the support and pelletizing.
A third modification consisted of soaking the support in a solu-
tion, or suspension, of catalyst and promoter, and evaporating
the solvent (water) to deposit the salts and oxide on the sup-
port.
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III. CATALYST TESTING PROCEDURE
A complete test of a given catalyst would Involve determining the
conversion obtained by varying each of the following:
Temperature
Concentration of SOz
Concentration of Oz
Contact Time
Concentration of other constituents>
active or inert
The rate of reaction is also dependent upon the catalyst1 surface: area,
pore volume, and size as well as the variables listed above.
Testing of the effects and interactions of all the variables on
conversion for a large number of catalysts would have required an
inordinate amount of time and expenditure for the amount and
usefulness of the information obtained. It was therefore decided
to pursue a more rapid but less comprehensive test program for
preliminary testing of catalysts.
Since the variation in SOz and Oz concentrations in stack gas is
not large and both are at low absolute concentration, average values
of their concentrations were employed and kept relatively constant
throughout the tests. Thus, a feed stream of 0.3% SOz and 3$ Oz
would be representative.
Next, a standard set of test conditions was established as follows:
Flow Rate - J»0 cc/sec (space velocity 2000-2500 hr"1)
"Standard" stack gas composition*
Initial run temperature - 600°F
It was then necessary to select a common basis for evaluating
different catalyst samples. A first approach considered "contact
time" as the common base. It soon became evident that this was
inadequate because the variation in bulk density, Table 1, among
different catalysts, resulted in widely different amounts of catalyst
showing the same contact time.
*Stack Gas Composition y , *
Nitrogen .......... 82.15
Carbon Dioxide. ... 14 . 70
Oxygen ............ 2. 80
Sulfur Dioxide.... 0.30
Nitrogen Oxides. . . 0.05
100.00
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Table 1
CATALYST BULK DENSITY
Catalyst
Code
A
B
C
D-l
D-2
E
F
G
Type
Vanadia
Pt
Vanadia
Vanadia
Vanadia
Vanadia
Pt
Pt/Fe
Weight Volume Apparent Density
(g) (cc) (g/cc) '
146.10
200.00
200.00
240
245
195
0.609
0.816
1.03
Sample Catalyst In Powder Form
Sample Catalyst In Powder F6rm
142.83 260 0.549
64.80 145 0.447
47.09 75 0.628
The 'Second approach considered the W/F value, i.e., the weight of
catalyst per mole of S02 per second. When results obtained using
contact time as a basis were converted to a W/F basis, as in Table
2, they were found to fit reasonably well with the computer-predicted
curves derived from rate equations in the literature as illustrated,
for example, in Figure 16. This, then, provided a common basis for
comparing different catalysts, namely: comparative conversion effi-
ciency at the same temperature and W/F value. Further, the W/F
value is related to space velocity (See Appendix y) thus permitting
reasonably valid extrapolation from laboratory test data to process
design.
In order to change the W/F value for a given weight of catalyst,
it was only required to change the gas flow rate. Later, conversion
isotherms are presented for some commercial catalysts at various
W/F values attained in this manner.
In summary, the procedure employed for the comparative evaluation of
experimental catalysts consisted of determining the conversion
efficiency (% S02 converted) at the same W/F value under the standard
test conditions enumerated above. In a number of instances, cata-
lysts were studied more broadly, as with some of the commercial
catalysts discussed below.
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Table 2
CATALYST TEST DATA
Weight of
Catalyst, W Flow, F
Catalyst (gms) (cc/sec)
1.
2.
3-
1.
5.
6.
7.
8.
9-
10.
11.
12.
13-
14.
15.
16.
17.
18.
19.
20.
21.
A
(pellets)
A
(pellets)
A
(pellets)
A
(pellets)
A
(pellets)
A
(crushed)
A
(crushed)
A
(crushed)
A
(crushed)
A
(crushed)
B
B
B
B
C
C
C
C
C
C
D-l
5.8
5.8
5.8
5.8
5-7
6.0
6.0
6.0
6.0
6.0
8.5
8.5
8.5
8.5
9.7
9-7
9-7
9.0
9.0
9.0
7,6
lo
60
60
60
60
40
40
60
60
60
60
kO
60
60
10
60
60
60
100
. 60
60
W/F
gm sec/cc
0.145
0.0966
0.0966
0.0966
0.0950
0.150
0.150
0.100
0.100
0.100
0.142
0.212
0.1*42
0.142
0.242
0.162
0.162
0.150
0.090
0.150
0.127
Temp.
°F
800
800
650
900
900
800
850
850
650
900
850
850
650
900
850
850
650
850
850
900
900
Percent
Conversion
38*
29*
3.5*
46*
25*
52*
68*
58*
3.5*
68*
50*
65*
15-11*
70*
62*
51*
8-3*
18*
23*
53*
57*
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Mori & Maenen 03
— —•———— Goldman «t. al.
Mars & Maeucn «1 425°C (800°F|
• Cotolyit A, cruihtd
1.0 2.0
W/F, gm-iac/mole SOjXlO"6
3.0
Figure 16
Comparison of Experimental and Predicted
Effects of W/F on Conversion at 800°F
26
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IV. CATALYST TEST RESULTS AND DISCUSSION
From statements found In the literature, there was reason to believe
that commercially available S02 oxidation catalysts, designed for
use in sulfuric acid plants, would be effective with flue gas. To
support this belief, commercial vanadia catalyst samples were ob-
tained from major U. S. producers. One platinum catalyst, commonly
employed in hydrocarbon reforming, was also obtained.
A. Commercial Vanadia Catalysts
In the particular case of vanadia catalysts, conversion efficiency
is rather poor below 790°F and above 930°F for a single pass con-
verter system due to greatly reduced activity at the lower tempera-
ture and the rapid shift in equilibrium at the higher temperature.
Consequently, three temperature levels were selected as representa-
tive of vanadium catalyst activity, namely: 800°, 850° and 900°F.
By experimentally defining the W/F conversion profiles at these
temperature levels, a rather complete picture of the relative merits
of the various samples was obtained. With the addition of catalyst
physical properties, it is possible to predict the weight of catalyst
needed for 90+# conversion of a given S02 concentration at constant
flow and temperature conditions.
Only two of the four commercial samples received were studied ex-
tensively. A few data points were obtained with the remaining two
to support the assumption that they too would, in general, effect
conversion of S02 under flue gas conditions.
Approximately 20 experimental runs were required to determine the
W/F conversion profiles presented for the two vanadia catalysts.
The conversion isotherms are shown in Figures .17 and 18. Examina-
tion of the curves indicates that 90$ conversion of S02 to S03 may
be attained at typical stack gas concentrations (^0.3$) providing
the indicated temperature, flow, and catalyst charge requirements
are met (See Appendix I for commercial catalyst test data).
As would be expected, there are point-to-point differences in the
performance of any two different commercial catalysts, shown in
Figures 19 through 21. This is not only true "between manufacturers"
but also "within manufacturers". Nevertheless, the data support the
conclusion that the commercially available vanadia catalysts are
applicable, in a technical sense, to a process for removal of sulfur
dioxide from flue gas. Furthermore, it is likely within the scope
of the manufacturers' present technology to adjust the performance
of these catalysts to give an incremental increase in efficiency
under flue gas conditions. Process economics and competition will
probably motivate such adjustment. Additional data relating to
vanadia catalysts can be found in Appendix III, Volume I.
27
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CATALYST -A-
900°F
850°F
— • 800°F
1.0 2D
W/F, gm.sec/mole SO2xlO"6
ao
Figure 17. W/F-Conversion Profiles for Catalyst 'A'
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100
Z
o
o
Z
iu
U
30-
20-
10
CATALYST E
900°F
850°F
800°F
1.0
I
2.0
3.0
W/F, gm.tec/mole SO,
Figure 18. W/F-Conversion Profiles for Catalyst-E
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100
90-
80-
70-
Q
ui
> 60
O
u
8 30-
40-
30-
20-
10
CATALYST A
CATALYST E
1.0
2.0
i
3.0
W/F. Om.««c/mol« SOj
Figure 19. Comparative Conversion Efficiency of Two
Commercial Vanadia Catalysts at 800°F
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100
90'
80-
70-
> 60-
O
u
8 50H
»-
z
iu
U
£ 40-1
30-
20-
10-
CATALYST A
CATALYST E
1.0
I
2.0
i
3.0
W/F, gm-i«c/mol* SO,
Figure 20. Comparative Conversion Efficiency of Two
Commercial Vanadia Catalysts at 850°P
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100!
10
CATALYST A
— — CATALYST E
1.0
i
2.0
I
3.0
W/F, gm-i«c/mol« SOj
Figure 21. Comparative Conversion Efficiency of Two
Commercial Vanadia Catalysts at 900°F
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B. Commercial Platinum Catalyst
Platinum catalyst designed for contact sulfuric acid plants is no
longer made commercially in this country. The platinum catalyst
sample studied here was designed for hydrocarbon reforming. How
much difference in composition or nature one might expect in
catalysts designed for these seemingly different purposes is not
known. However, as indicated by the conversion isotherms in
Figure 22, this particular catalyst .gave a creditable performance.
Furthermore, a point-to-point comparison of the performance of
the platinum and vanadia catalysts at flue ;gas conditions presents
essentially the same pattern as that found in the literature for
the comparative performances of the two types -of catalyst at
"standard" contact plant conditions. This can be seen in Figure
23 for comparative performance, at 900°F.
In contrast to the vanadia catalyst, the platinum displayed an
unstable equilibrium at constant reaction conditions for a1con-
siderable period of time at. each set of conditions .tested. The
effect of time on the conversion efficiency is shown in Figures
24 and 25: Sometimes as long as six hours was required before
a stable, reproducible conversion was obtained.
Although vanadia and platinum .catalysts would seem to be about
equivalent in performance, they are worlds .apart in cost. A cost
analysis of these two catalysts, summarized in Table 3, indicates
that platinum must be at least ten times as effective as vanadia
to be comparable in cost with the vanadia catalyst. Decline in
platinum catalyst cost is not likely in the near future in view
of the rising costs of labor and the metal .itself .(Details in Appendix VI
C. Experimental Catalysts.
Considering only the performance, characteristics of vanadia and
platinum catalysts, a major disadvantage with both, relative to
a flue gas process, is their high reaction temperature. This
generates both economic and engineering difficulties in practical
application. Consequently, it seemed logical to seek a catalyst
with an operating temperature nearer to thstt of the flue gas enter-
ing the stack, i.e., about 300°Fi This temperature, however, is
well below the dew point of the sulfuric acid produced through
oxidation of the flue gas sulfur dioxide. Thus,, a catalyst•system
operating at this temperature would soon be swamped with liquid
sulfuric acid.
A corollary to the "low" temperature approach is identification of
a catalyst which is more active throughout the present operating
temperature range than are commercial vanadia catalysts. If a
33
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— 900°F
.... 850°F
_ 800°F
i
1.0
I
2.0
i
3.0
W/F, gm.i«c/moU SOj
Figure 22. W/F-Conversion Profiles for Commercial
Platinum Catalyst
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lOOi
10
PLATINUM (CATALYST B)
VANADIUM (CATALYST A)
T
1.0
T
2.0
3.0
W/F, gm.8«c/moU SO} *10~
Figure 23. Comparative Conversion Efficiency of
Platinum and Vanadia Catalysts at 900°F
35
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40,
36.
32
28-
at
in
> 24
Z
o
u
£201
t-
z
ui
U
S 16.
12-
4 .
CATALYST B
20 40 60
ELAPSED TIME AT REACTION CONDITIONS, minutei
Figure
Effect of Time on Conversion Efficiency of
Catalyst-B at 800°F, Constant Reactor
Conditions (Run No. 6)
36
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Q
IU
Z
o
u
65
60
55
50-
45-
O
to 401
at
a.
30-
25-
20
15
CATALYST B
1
30
I
60
i
90
ELAPSED TIME AT REACTION CONDITIONS, minutes
Figure 25. Effect of Time on Conversion Efficiency
of Catalyst-B at 800°F, Constant Reactor
Conditions (Run No. 7)
37
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Table 3
SUMMARY OF ECONOMICS FOR VANADIA AND PLATINUM CATALYSTS^
V Catalyst Ft Catalyst
Volume, Cu ft^1) 30,300 3,020
Initial cost, dollars^2) 1,250,000 12,800,000
Regeneration cost, dollarsC2) 1,032,000
Capitalized cost, dollars(2) 3,920,000 20,380,000
NOTES: (1) Estimate of the volume of platinum catalyst that
would be equal In cost to the estimated capitalized cost of
vanadia catalyst for 90$ conversion in flue gas.
(2) Estimates are based on the same W/F ratio for
platinum and vanadia catalysts.
(3) Assumptions used for comparison of platinum and
vanadia catalysts:
General:
a. rate of return: Q% per year
Vanadia Catalyst:
b. amount used for the oxidation of SOz in flue gas:
30,300 ft3
c. price: $4l/ft3
d. packed density: 36.8'lb/ft3
e. catalyst life: 5 years
Platinum Catalyst:
f. catalyst contains Q.5% platinum
g. packed density: 50 lb/ft3
h. catalyst cost: cost of platinum + $2.75/lb
catalyst, (for manufacturing cost)
i. regeneration cost: $0.75/lb catalyst
j. regeneration interval: 2 years
k. loss of platinum during regeneration: 2%
1. price of platinum: $110/oz (troy)
m. catalyst life: 30 years (including obsolescence
and/or abandonment)
38
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catalyst were, say, 25-30$ more effective than vanadla throughout
the activity range, then at a temperature of 750°F, where vanadia
shows only 60-705& efficiency, the new catalyst may yield a desirable
90$ conversion efficiency.
Yet a third objective in testing experimental catalysts was to follow-
up suggestions, or clues, in the literature that could lead to new
catalytic compositions for efficient application to flue gas clean-
ing.
Observations made during the course of the commercial catalyst evalu-
ation studies appeared to shed some light on the mechanism of
vanadia catalysis as proposed in the literature:
S02 + 2V+5 + O-2 * S03 + 2V+" (1)
2V+1+ + 1/2 02 f 2V+5 + O-2 (rate con- (2)
trolling)
V02, V205 (solid) *V02, V205 (in solu- (3)
tion in molten alkali
pyrosulfates)
In our catalyst test unit, at the end of a run, the mixed gas
stream is turned off and replaced by a stream of nitrogen at the
same bed temperature. The nitrogen purges the system for roughly
an hour at temperature; then the unit is shut down. This opera-
tion, supplanting the mixed gas stream with nitrogen, has the ef-
fect of "freezing" the equilibrium mixture in Equation (2). We
compared pellets of a catalyst "frozen", in this manner, at 625°F
and at 800°F by placing them separately in distilled water. The
pellets representative of the lower temperature were bluish in
color and produced a blue solution, characteristic of V+1*. Pel-
lets representative of the higher temperature were brown, charac-
teristic of V+5 and imparted no color to the water. This qualita-
tive observation appeared to confirm the belief that at the lower
temperature, certainly the equilibrium mixture was predominantly
V+4*, but it raised the question of whether Equation (2) is inherently
temperature-dependent or whether it must act in molten solution,
which condition is temperature-dependent. If it were solely the
latter, then a possibility of defining a system of low melting
promoters is offered.
There are salts which melt in the desired "low" temperature range,
and using some of these, fleeting evidence supporting the rationale
above was observed. The short life of activity in the lower tempera-
ture region is attributed to instability of the salt either at
temperature or in the presence of S03.
39
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Again we have seen vagrant activity of this type when, for instance,
residual chlorine in platinum catalyst is replaced with thiocyanate.
The increased activity of the platinum was short-lived because,it
Is proposed, of the reaction of sulfuric acid produced with the
thiocyanate to release thiocyanic acid.
Many catalyst compositions were prepared and tested in the search
for a stable low temperature, or more active high temperature
catalyst. Included also were a number of experimental catalysts
from various vendors around the country. These data are quite
extensive and, for that reason, are tabulated in Appendix II. All
catalyst compositions beginning with the code letter (M) are
compositions prepared in our laboratory. Other codes represent
experimental samples received from industrial catalyst vendors.
The latter were not accompanied by details of their compositions,
as this information is considered proprietary by the vendors.
It is interesting to note, throughout Appendix II, several com-
positions (designated by stars) which exhibit extremely low
temperature activity. The apparent conversion efficiency reported,
while accurate, may be somewhat deceiving, as these data may not
represent optimum operating conditions of the catalyst.. However,
the data does indicate the possibility that these compounds have
potential as stolchiometric removal .systems similar to the limestone
injection process. As mentioned previously, their performance may
be affected by several parameters not.yet defined.
It is felt, that the behavior nearly always observed, i.e., fleeting
low temperature activity, can be explained on the basis of the
product SOa reacting with the metal oxide catalyst forming sulfate,
or with the promoter, as with KHSOi|:
2 KHSO,, -I- S03 »- K2S207 + E2SOk
As the bisulfate is converted to pyrosulfate, the melting point of
the salt rises, with a corresponding decrease in activity at the
lower temperature of the test.
These observations are interesting in that they seem to support our
approach to a low temperature catalyst. But, they also illustrate
the necessity of stability of the promoter system in the presence
of 863. No material tested showed greater activity or stability
than vanadia under test conditions.
In the light of these cumulative observations, it is possible to
prescribe the requirements for a "low" temperature vanadia catalyst,
which, in the end, resolve to the requirements for the promoter
salt.
40
• MONSANTO RESEARCH CORPORATION •
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These requirements are:
a. Salt (or salt system) should be completely molten
and non-volatile in the 450°-500°F range.
b. Vanadia must .dissolve in molten salt and, probably,
must, ionize.
c. Salt (or system) must be chemically stable in tempera-
ture range.
d. Salt (or system) must be chemically stable in the
presence of 02, S02 and S03.
These requirements assume that sorption and desorption of reactants
and products are not limiting, which is the current assumption for
commercial vanadia catalysts.
D. Molecular Sieve Adsorption Studies
Molecular sieves had been shown in the literature to be useful in
many applications as heterogeneous catalysts and as vehicles in
studies of catalysts and catalyst mechanisms. However, we could
find no publication which revealed specific studies to determine
their merit as support material for S02 oxidation catalysts. As
the monodisperse nature of the zeolite pores and their extremely
high specific area were characteristics favorable for maximum
contact conditions between the S02 and active catalyst sites,
samples were obtained and used as support material for two previously
tested compositions. No alteration of the catalyst characteristics
was seen.
Zeolite materials are also well known for their adsorption character-
istics with regard to the various constituents in stack gas, and we
wished to know if these characteristics would be an aid or
hinderance in their use as a support material. Adsorption efficiency
profiles at four temperature levels between ambient and 250°F were
established for (4) different molecular sieve materials. These
data are shown in Appendix VII. " As regards materials
SK-410 and SK-400, both of which showed relatively high S02
adsorption efficiency, the effects of a 30$ reduction in adsorption
bed contact time are 'also shown.
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E. SO? Removal Characteristics, of "Red Mud"
During the course of our experimental catalyst testing program we
were requested by NAPCA to perform evaluations of a material
commonly termed "red mud," which we were informed is a waste by-
product of aluminum ore processing operations, for its potential
as a removal material for S02 in power plant stack gas.
Using the one sample sent to us, we ran a bed of the material
through our catalyst testing system and obtained conversion efficiency
vs. time profile as shown in Figure 26. In a further attempt to
determine the constituents influencing the activity of red mud in
the presence of SC>2, a mass spectrometer analysis was performed on
a sample of exhaust gas from the reactor during the time period
shown in Figure 26. No evidence of S03 was found. However, these
results are inconclusive, as we were unable to determine the
sensitivity of the instrument to small concentrations of 803'.
Further tests were performed on the partially spent bed of red mud
which produced the decay profile illustrated in Figure 26. Emission
analysis of the sample revealed the presence of the, following:
Fe — Low Major
Si — Low Major
Al — Low Major
Ca — Low Major (^10/S)
Na — 5 to 10%
Ti ~ 5 to 101
Zr — 0.3%
Cr — 0.1%
V —• 0.07?
Mg — 0.03*
Mn — 0.03?
Subsequently, duplicate fusions were performed on the spent sample
to determine total sulfur content. Results were 5-38 and 5-3^%
sulfur. A water and acid insoluble residue resulted when the
fusion was taken up in water. The best probability is that the
residue was a titanium containing material.
According to stepwise numerical integration of the absorption
versus time profile (Figure26) the sample of red mud absorbed 288 cc
S02 over a 95 minute contact time. The density of SC>2 at metering
conditions is 2.6813 g/1. Accordingly, the sample absorbed 0.772
gm S02. Of this, 50% of the total would be sulfur, or 0.386 gms.
The sample weighed 9.82 gms, and of this an average 5-36/? was
determined as sulfur, or 0.526 gms. Consequently, 0.140 gms of
sulfur were unaccounted for.
• MONSANTO RESEARCH CORPORATION •
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z
o
•z.
w
z
H
O
X
m
en
m
> -t
a LO
o
I
o
o
33
6
z
90
80
g 70
o 60
LU
OH
O
o
50
o
on 40
30
20
Analysis Sample Taken (Mass Spec.)
~K) 20 30 4& 50 60 70"
ELAPSED TIME, minutes
80
90
100
Figure 26. Conversion Efficiency versus Time Profile & Constant Operating
Conditions for "Red Mud"
-------
Duplicate fusions were performed on a fresh sample of red mud
with a result of Q.J2% sulfur, or 0.071 grts sulfur In the sample
Initially. Thus, half of the missing sulfur was accounted for.
This left about 13% of the total sulfur unaccounted for. However,
this Is within the accuracy of the overall technique.
These results Indicate that the activity of the red mud is due
to reaction between the S02 and various constituents of the sample,
with subsequent formation of stable inactive sulfur compounds
within the sample, causing the decrease in "apparent conversion"
noted in Figure 26.
• MONSANTO RESEARCH CORPORATION •
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V. DEVELOPMENT OF A SIMULTANEOUS SOX-NOX REMOVAL PROCESS
It has long been recognized that high temperature combustion processes,
including electric power plants, produce oxides of nitrogen. It
was therefore a logical and desirable extension of our experimental
program to check the effect various commercial catalysts had on
existing NOX concentrations in a power plant stack gas environment.
In addition, we felt that a process which would simultaneously
remove both SOX and NOX would have a decided advantage over a
process removing only one of the pollutants.
Examination of the well-known technology of the Chamber Sulfuric
Acid process indicated that nitric oxides could be scrubbed from
their carrier gas by 60° Be sulfuric acid - successful removal being
dependent upon the maintenance of equimolar concentrations of NO and
N02• The resulting absorption of NO + N02 into sulfuric acid then
produces nitrosylsulfuric acid (NOHSOi,), this formation resulting
from the following reactions:
• NO + N02 + 2H2S04 >- 2NOHSOI, + H20 (1)
• 2N02 + S02 + H2S04 >- 2NOHS04 (2)
• 2N02 + H2SOt» >- NOHSOit + HN03 (3)
The nitric acid which accompanies the formation of NOHSOi* In equation
(3) is then soluble in the large excess of sulfuric acid.
A. Initial Laboratory Investigations
Realizing that a similar chemically favorable environment could be
produced.in the exhaust stream from our catalytic reactor, a small
glass scrubber was constructed and installed downstream of the reactor
as represented in Figure 27« Sufficient N02 was then added to the
flue gas to meet the following minimum requirements:
• 2 mols N02/mole unreacted S02 in flue gas
• 1 mole N02/mole NO in flue gas
As we had no means available at this time to meter the amount of N02
into the system, an obvious excess was used (as evidenced by the
characteristic color of N02 in the combined gas streams). The
rationale here was that the excess N02 would, react according to
equation (3)•
As soon as N02 was added to the reactor effluent, white crystals
formed on the walls of the glass arm entering the bottom of the
scrubber. The accumulation of these crystals shortly plugged the
45
• MONSANTO RESEARCH CORPORATION •
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60°Be' I H2S04
ACID
DISTRIBUTOR
CONVERTER
EFFLUENT
CHAMBER
CRYSTALS
TO
CHROMATOGRAPH
ACID OUT
Figure 21. Laboratory Arrangement for NOX Scrubber
• MONSANTO RESEARCH CORPORATION •
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line. These crystals were qualitatively identified as "chamber
crystals" or the anhydride of nitrosylsulfuric acid. This was not
surprising, as the gas mixtures were dry. More interesting, however,
was the reduction of unreacted SC>2 in the flue gas. At the time,
the catalytic converter was effecting about 90$ S02 conversion,
but the exhaust from the scrubber showed 95+% conversion.
The NC>2 feed point was moved to the scrubber and the experiment
repeated with no apparent formation of chamber crystals. The
exhaust from the scrubber was then analyzed to determine the effect
of NC>2 on the total conversion of sulfur dioxide. These data are
shown in Table 4.
Table 4
EFFECT OF N02 ON TOTAL CONVERSION OF SULFUR DIOXIDE
Momentary total conversion without N02
addition to scrubber 89.1/8
Momentary total conversion with N02
addition to scrubber 1) 9^.9$
2) 100.0$
3) 99.2$
4) 99.2$
v
Resulting product acid was shown qualitatively to contain nitro-
sylsulfuric acid.
The surprisingly high total conversion of S02 resulting from the use
of N02 in the scrubber prompted further experiments. The flue gas
was fed directly into the scrubber, by-passing the catalytic
converter. Results of this experiment and a comparison with and
without the use of the converter are given below.
Flow Scheme % S02 Conversion
Source -»• Converter -»• Scrubber 89.1
Source -»• Converter •*• Add N02 •*• Scrubber 99-5
Source -»• Add N02 -»• Scrubber 99.7
47
• MONSANTO RESEARCH CORPORATION •
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B . Vapor Pressure Studies on the System; H2 SOU -H2 0-N2 0 3
Our success in obtaining "almost 1005? -SO 2~ oxidation by addition -of
N02 prompted the decision to design a process based on "modified"
Chamber technology. Initial attempts to design such a process
pointed out the need for vapor pressure data on the system:
^SOit-^O-^Oa . A literature search revealed two articles on the
subject - one by Berl (Ref. 1) in 1934 and another by Tseitlin and
Yavorskii (Ref. 2 ). Berl's data covered the range of sulfuric
acid concentrations from 64.05? to 80.05?, at temperatures from 30°C
to 150°C and for N203 concentrations from 0 to about 1%. The data
of Tseitlin and Yavorskii covered an acid concentration range from
81.2% to 84.14$ for corresponding N203 concentrations from 4.8% to
about 10.05? at 50°C, 70°C, and 90°C.
A brief vapor pressure study was also initiated to check the data
of these authors and to extend their work over the acid concentration
range of 64 to
A pictorial representation of the experimental apparatus is given in
Figure 28, which, incidentally, is quite similar to Berl's as given
in Appendix 9. A 100-ml. round-bottom flask was filled approximately
half-way with the desired ^SO^-^Oa mixture. The flask contained
a teflon coated magnetic stirring bar. The flask was immersed in an
oil bath with temperature maintained, by an electric heater. A (-10°
to 200°C) thermometer was suspended in the bath for temperature
measurement. The flask was connected to one of two Hg manometers
by means of rubber and glass tubing - the choice of manometer
depending upon the expected vapor pressure. A MacLeod gauge was
used to measure pressures from about 0.01 mm Hg to 5 mm Hg, and a
U-tube manometer for pressures between 3-700 mm Hg. The system
was connected to a vacuum source equipped with a liquid nitrogen
trap and a dry ice trap which could pull a vacuum down to 0.007 mm
Hg. There were two pinch clamps on the rubber tubing - one of which
was immediately above the round bottom flask, and a second on the
line to the vacuum source.
Our initial experimental results indicated that the data of these
authors was low by a factor of about 3« However, we found that our
procedure used to determine the vapor pressure was faulty. When this
fault was corrected, the data agree satisfactorily with the data of
Berl and Tseitlin and Yavorskii.
A description of the apparatus and procedure used to obtain the
vapor pressure data is given in Appendix 8 to point out the faulty
procedure and thereby prevent its recurrence in the future. The data
is not included because it is of rather poor quality compared to
that of Berl, Tseitlin, and Yavorskii, and would only serve to
48
• MONSANTO RESEARCH CORPORATION •
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o
z
z
o
m
to
PI
>
TO
O
I
o
o
3J
TJ
O
o
z
Heater & Controller,
Glass Tubing
100ml Flask
Stopcocks
Rubber Tubing & Pinch Clamps
Hg
Manometer
—To Vacuum&Liquid Na&Dry IceTrapsI
Magnetic Stirrer
Figure 28. Experimental Apparatus for Vapor Pressure Studies
-------
confuse the picture. Instead, a summary of their data is given in
Figure 29- Berl's data was plotted on a log P(NO and NOa) vs 1/T
plot for 1.0$ N203 mixture, and was limited to the following acid
concentrations: 64.0$, 67.0$, 68.5$, 70.0$, 73.3$, 75.65$, 78.0$,
and 80.0$. The best line was drawn through each of the acid
concentrations. As the lines appeared to intersect near a single
point, the "best point" was chosen and all lines were drawn to
Intersect at that point. The lines at 71 and 72$ were included
for convenience. Lines above 80$ were obtained by extrapolating
Tseitlin's and Yavorskii's data back to 1$ ^03 and forcing the line
to pass through the intersection point. The greatest deviation
from experimental data appears to occur at about 80$; however, these
deviations differed in direction for the two authors. Berl's data
shows somewhat lower NO and NC>2 vapor, pressures (at 110°C and up)
than that shown on the plot, whereas Tseitlin's and Yavorskii's data
show somewhat higher NO and N02 pressures at 80$.
The vapor pressure data given in Figure 29 can be extended to other
N2©3 concentrations by multiplying by the ^03 concentration in
wt$.
The design calculations given in this report were based on the vapor
pressure data given in Figure 29.
C. Absorption Studies of NO and NOg into Sulfuric Acid
A laboratory absorption column was constructed to test the kinetics
of absorption for NO and N0£ into H2SOi» and to compare experimental
absorption efficiencies with predicted efficiencies using vapor
pressure data. Figure 30 is a schematic representation of the
column. Two gas cylinders were used, one containing nitrogen and
the other NO, N02 and N2. The amount of gas fed from each cylinder
was adjusted to provide NO and N02 vapor pressures of about 3 mm Hg
each. These gases then passed into a mixing chamber to insure that
the gas fed to the column would be homogeneous. The gas then
passed through an electrically heated tube and into the bottom of
the column. The column and acid reservoir were heated by exterior
heating tapes, and insulated with 1/2 in. - 1 in. Pyrex wool. A
description of the procedure used in this study is given in Appendix 9.
Table 5 is a summary of the data obtained from this study. Eight
runs were made - the first three using a column packed with 1/2 in.
Raschig rings, and the last five with a column packed with cut Pyrex
tubing. This tubing consisted of about equal portions of 5 mm and
6 mm tubing cut to 5 mm and 6 mm lengths respectively.
Examination of the data revealed that the efficiency of the 1/2 in.
Raschig rings to remove NO + NOz from the gas was about an order of
less than the cut Pyrex tubing.
50
• MONSANTO RESEARCH CORPORATION •
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1000
0.1
300 280 260 240 220 200 180 160 140
120
100
80
40
c (-fr SCALE)
Figure 29- Adjusted Vapor Pressure of NO and N02 Over Nitrose
• MONSANTO RESEARCH CORPORATION •
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Acid
Storage
Acid Control Valve —
2.5' X 1.9'
Column
Electric
Heaters
Thermometer
(
Gas Sampler
FoVent
Wet Test Meter
Thermometer
Gas
Sampler
LJ
Cylinder
SCZZKI
Mixing Chamber
Rotameters
Metering Valves
N,-NO -NO,
Cylinder
Figure 30. Absorption Studies Apparatus
• MONSANTO RESEARCH CORPORATION •
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Table 5
SUMMARY OF LABORATORY N203 ABSORPTION STUDIES
Estimated
NO + N02
z
0
z
z
H
o
3
rn
en
m .„
^
o
i
0
o
TJ
0
TO
H
Press
Inlet
1.10
0.14
3.8
3.8
3.8
8.8
3.78
4.82
(mmHg)
Outlet
1.00
0.138
3.4
1.2
2.1
4.3
1.41
0.74
NO + N02
Removal
( %}
9-1
1.0
10.5
68.4
44.8
52.3
62.7
85.0
Approx.
Flow
Gas
cc/Min
12000
12000
12000
6000
6000
12000
12000
12000
Rate
Acid
cc/Min
15
80
100
80
33
35
55
60
Temperatures (°C)
Gas
In
119
120
126
26
26
82
25
110
Gas
Out
117
126
122
26
26
75
25
110
Acid Acid
In Out
129 75+
110
112
26
-------
In an attempt to define the cause of the poor NOX absorption, a gas
sample of the N2-NO-N02 mixture was taken and submitted to I.E.
analysis. The analysis indicated that the sample had a respective
gas ratio of 2:1:0.4 as N02/NO/N20. This unexpected appearance of
N20 in low concentration was found to be a result of using a very
low grade NO supply. The NO cylinder used to blend the mixture
contained about 25% N20 and about 10$ N02. Thus, of the total NO
and N02 present, only 50$ existed as a 1:1 ratio of NO to N02-
Since the removal of NO + N02 from flue gas is optimum with a 1:1
ratio of NO to N02, and falls off sharply as the ratio differs, it
is not surprising that the average removal of NO + N02 was about
63$ (avg. of last 5 runs).
Although the absorption studies did not show the desired 99-6$ NOX
removal, they did point out the necessity of. maintaining a'lrl mole
ratio of NO to N02. If this ratio is not maintained closely, the
recovery of NO + N02 will be severely reduced.
D. Proposed SOX-NOX (SONOX) Removal Process Description
1. General Description
Utilizing the data.gathered during our laboratory investigation of
the dual purpose removal system, a process design was developed as
shown in Figure 31. In this process, flue gas enters the bottom of
a trifunctional absorption tower. In the bottom section flue gas
is dewatered to ^0.2$ to 0.5$ H2.0 and cooled to 1?0°F by counter-
current contact with 160°F, 84$ H2SOi+. The acid is then diluted
from 84$ to about 8l$ and heated to about 290°F. The cooled flue
gas then enters the middle section of the tower where recycle. N02
is added. The N02 oxidizes S02 to SOa and the resulting NO reacts
with residual N02 to form ^03. The heat of reaction raises the gas
temperature about 20°F while the 803 and N203 are absorbed in, 80$
f^SOi* at 190°F. The flue gas then enters the third section of the
tower where it contacts the acid removed from the bottom section.
The gas becomes humidified and heated while the acid is reconcentrated
to 84$ and cooled to about l60+°F. The acid may have to pass through
a cooler to bring the temperature back to l60°F, unless there is
evaporative cooling.
The acid removed from the middle section passes through a heat
exchanger and into a stripping tower. Here the, dissolved ^03 is
stripped from the acid at 338°F with hot gas. The denitrated acid
leaves the stripper, passes through the heat exchanger, a cooler,
and back to the middle section of the absorption tower. A small
product stream is then removed beyond the heat exchanger. The
stripper exit gas enters an oxidation chamber which provides
time to effect 80+$ conversion of NO to N02.
54
• MONSANTO RESEARCH CORPORATION •
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2
O
z
u>
>
z
H
O
m
W
m
o
i
o
o
7)
•0
O
O
Z
U1
25% Flue Gas to each of
Four Towers
Flue Gas
300° F
7100 moles/min
7.25% *
H20
Clean Flue Gas
270° F
50 moles/min
Product H SO
Product N02
-3.5 moles/min
Recycle N02
-50 moles/min
-10% NO,
122°p
Cooling Water
80%* 900
moles
(N02)
(-12.5 moles/min)
40psig Steam
50moles/min
3600 moles/min
81% Acid-290° F
lOOOmoles/min
Condensate
80 moles/min
Vaporized
Condensate
Product Acid Out
~50moled/min
ONE OF FOUR TOWERS
ONE STRIPPER
Figure 31. Process Design Schematic
-------
Water vapor is condensed by cooling the gas to 122°P with water.
The NOa-rich gas stream, minus a small product stream, is then
recycled to the middle section of the tower. The condensate from
the oxidizer is revaporized and fed back to the stripper. In
addition, make-up water is added to the stripper as steam.
2. Trifunctional Absorption Tower
a. Bottom Section of Absorption Tower
Vapor pressure studies on the system I^SO^-HaO-^Oa revealed the
desirability of operating the absorption section of the tower, at
as low a temperature as possible and with an acid concentration of
80+/K. In order to meet this requirement, the flue gas must be cooled
to a suitable temperature (^170°F) and must be dewatered to the
equilibrium composition of H20 vapor over 80% sulfuric acid at ^190°?.
The requirement that the gas must be dry arises from the fact that
absorption of H20 into the acid would be accompanied by a decrease
in acid concentration and a rise in .-temperature from the heat of
absorption - both detrimental to NO + N02 absorption.
The following assumptions were made in designing this section of the
tower:
1.36 pounds of 84$ HaSOit/pound of flue gas
The temperature of the acid leaving the tower will be
10°F cooler than the entering flue gas temperature
b. Middle Section of Absorption Tower
In this section the S02 in the flue gas is oxidized to SOa by NC>2
(obtained from the recycle stream), thereby permitting removal of
the sulfur as H2SOi+. The amount of N02 recycled is controlled to
affect a 1:1 ratio of NO to N02 after all the S02 is oxidized. The
temperature of the recycle N02 stream is about 122°P and the heat
of S02 reaction is sufficient to affect a 20°F rise in the flue gas
temperature. Eighty (80) percent H2SOit is used to scrub the NO + N02
gases from the flue gas. At these operating conditions, approximately
0.88 pounds of acid/pound of gas will remove 99.6!? of the NO + N02
with 15 theoretical trays. The flue gas leaving this section
contains about 0.25% H20, 0.003? S02 , and about 0.005% NOX, and the
acid leaving this section contains about 1.08 wt# N203 (3-92$ HNS05).
Here, the following assumptions were made:
The oxidation of S02 to S03 is instantaneous
Fifteen theoretical trays are desired
No oxidation of NO takes place in the absorber.
56
• MONSANTO RESEARCH CORPORATION •
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c. Top Section of Absorption Tower
The function of this section is to utilize the dry clean flue gas
to concentrate sulfuric acid from Bl% to 84$. As the Ql% acid
enters the top of the section (from the bottom of the bottom section)
at 690°F, the gas passes countercurrent to the acid and removes
H2
-------
5. NO Oxldizer
The function of the NO oxidizer is to provide sufficient residence
time to allow the oxidation of NO to take place. Only one tower
(about 35' x 30' diameter) is required. This provides 9 sec.
contact time.
Assumptions made in this unit:
80+/? conversion of NO is desired
The gas is cooled to 122°F
The reaction constant is 1.5 x 101* (l2/mole2- sec).
6. Acid-Acid Heat Exchanger
The function of the acid-acid heat exchanger is obvious. The size
of the heat exchanger (surface area) depends upon the heat transfer
coefficient and the AT. For an assumed heat transfer coefficient of
150 BTU/hr'ft2-°F and a AT of 14°F, the required surface area is
3^5,000 ft2.
Assumptions made for acid-acid heat exchanger:
Temperature driving force is 14°F
• Heat transfer coefficient is 15 BTU/hr-ft2•°P.
7. Heat and Material Balances
The material balance for this process is given in Table 6 . .The
flow rate of each compound in a given stream is shown in lb-moles/-
min. The temperature and pressure of each stream is included if
significant.
E. Comparison Between Tyco and Monsanto Modified Chamber Process
1. General Description of Process Flow Differences
The description of the Tyco modified chamber process was taken directly
from a report issued by Tyco representatives at the Second Annual
Control Process Contractors' Meeting (NAPCA) held on June 11-13, 1969,
in Cincinnati, Ohio. A copy of that report is included in Appendix VIII
The Tyco process basicly rests on two important design features:
(1) the absorption of SOa and N203 (NO + N02) is accomplished in an
absorption tower operating isothermally at 250°F using 80/J I^SOit,
',_) tne absorbed N20a (HNSOs) is oxidized in the liquid phase over
a charcoal catalyst and the resulting N02 is removed by the gas
sweeping through the column.
58
• MONSANTO RESEARCH CORPORATION •
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Table 6
HEAT AND MATERIAL BALANCE - HIGH TEMPERATURE STRIPPER
*
z
O
•z.
3>
•z.
H
O
a
m
m
> VJ1
a vo
i
o
0
TJ
0
j>
H
•z.
0
Stream
Flue Gas
Gas Exit
BTM
Gas ABS
Exit
Clean
Flue
Recycle
N02
Prod. NO 2
Oxldlzer
Eff.
Stripper
Gas Out
Strip Air
ABS BTM
Acid Out
ABS Top
Acid Out
ABS Feed
Acid
ABS Acid
Out
Strip
Acid Out
Prod. Acid
Cond.
Oxldizer
H20
Make-Up
H20
N2
5318
5318
5578
5578
260
16.5
276.5
276.5
276.5
__
__
__
_
_ _
,„
—
03 C02
199 1044
199 1044
258 1044
258 ion
59
3-7 —
62.8
73-5 —
73-5 —
— —
— —
.
— —
— —
-- —
__
—
H20 S02 N02 NO H2SOU N203
515 21.3 0.35 3.20 —
21 21.3 0.35 3.20 —
21 0.3 0.17 0.17 —
515 0.3 0.17 0.17 —
28.9 — 45.45 5-05 —
1.8 — 2.88 0.32 —
30.7 -- 48.33 5-37 —
112.3 — 26.85 26.85 —
—
3106 — — — 2483
2612 — — — 2433
2066 — — — 1504 0.08
2075 — ~ — 1525 26.93
2094 -- — — 1525 0.08
28.8 — — — 21.0
81.6
51.6 — _
Temp.
(°F)
300
170
190
270
122
122
122
338
565
290
160V
100
203/
190
190/
323
338/
203
203
122/
380
380
Press.
(mm Hg) Comments
785
780
775
770
790
790
790 Remove 2.86xl06 Btu/min
800
•\-900
__
Cooler may be required
—
—
__
40 pslg
} ^.$800,000/yr
40 pslg
-------
Of the two features the second Is more important. It makes it
possible to operate the stripper at the same temperature as the
absorber. Thus, there is no necessity to heat the acid in order to
raise the NzOa vapor pressure and permit desorption, and later cool
it to the absorber temperature. The rest of the process consists of
(1) recovering nitric and sulfuric acids, and (2) preparing the flue
gas to enter the absorption tower by cooling to 250°F and reacting
NOa with S02.
The Monsanto process differs from the Tyco process on both important
design features. The absorption tower is trifunctional rather than
monofunctional. In the first section of the tower the gas is cooled
to about 170°F and dried to approximately 0.2-0.5? HaO with 84$
sulfuric acid. In the second section, N02 is added, oxidizing the
S02 to S03 and the 80s plus NO and N02 are absorbed with 80$ H2SOi+
at 190°F. In the third section the flue gas is warmed and moistened
by passing through the hot, humid acid removed from the bottom of
the first section.
The stripper is also different. Assuming that neither charcoal or
any other catalyst will effect the liquid phase oxidation of NO,
the recovery of NO + N02 is accomplished by raising the acid
temperature from 190°F to 338°F. At this, temperature, the vapor.
pressure of NO + N02 is about 30 times higher than it was at 190°F,
the absorbing temperature. Thus, the stripping of NO = NOa should
be accomplished with less than 5% of the flue gas volume.
2. Detailed Process Comparison
Table 7 is a direct comparison between the Tyco and Monsanto,processes
The comparison is made as the two processes were described. Tyco's
stripper is catalytic and Monsanto's stripper is high temperature.
The required absorption acid flow rates were calculated using the
vapor pressure data shown in Figure 29 as a common basis. The major
differences between the two processes are the following:
Tyco process requires about twice the acid flow rate in
the absorber as Monsanto's process
Monsanto process is circulating 1.08 times as much acid
as Tyco process
Monsanto absorption towers will probably cost 2 to 2.5
times as much as Tyco towers
Monsanto absorption tower is about 3 times higher than
Tyco towers
Monsanto process uses 24,000 GPM cooling water while Tyco
uses 17,000 GPM.
60
• MONSANTO RESEARCH CORPORATION •
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Table 7
COMPARISON OF MONSANTO AND TyCO PROCESSES (AS DESIGNED)
Subject
Monsanto
Gaa'Cooling
9.
10.
11.
12.
13.
ID.
15.
16.
Product Acid Concentra-
tion
Relative required add
flow rate for NiOa
absorption
Stripper Operation
Stack Oas Exit
Temperature
NO Oxidation Chamber
Total Relative Acid
Circulation Rate
(Absorber Stripper)
S02 Reactor
Acid Heat Exchangers
Required
Acid Coolers
Maximum HNSOs cone.
In absorber (at
minimum L/C)
Relative Absorber Tower
Construction Difficulty
Stripping Tower Con-
struction Difficulty
Catalyst Losses
Pressure Drop by Flue
Oas
Cooling Water Require-
ments
Uses a flue gas cooler to remove
1.3xl06 Btu/min as low-grade steam
Area of heat transfer • 90,000
ft2 (cools gas to ^278°P)
80S
1.0 (^1814,000 Ibs/mln)
Uses temperature increases to strip
N0+N02. Air Is used as stripping
gas (possibly with some flue gas).
No catalyst Is assumed. Required
gas rate Is 350 moles (at least
120 must be air)
270+°P
A tower 30* 0 by 36" (provides
•>-9 sec contact time)
1.08 (have three separate
absorption sections)
Done In absorber
One acld-to-acid exchanger with
327,000 ft2
8 (total 15,000 ft2)
6.08
2-2.5
1.0
None
•••26" HjO
24,000 GPM
Tyco
Cools all flue gas to 250°F;
removes heat of SOj oxidation;
removes heat of NO oxidation
Heat removal =
2.9xl06 Btu/min - cooling
O.SxlO6 Btu/min - SOz ox.
0.6xl06 Btu/min - NO ox.
80S (limited to 92* maximum)
2.0 (•v.381,000 Ibs/mln)
Uses charcoal as a catalyst to
oxidize N203 to 2 N02. Flue
gas Is used to oxidize NO and
strip N02 from acid. Requires
minimum of 850 moles/mln of
flue gas to provlda sufficient
02 for oxidation (if no air Is
used)
None, required If charcoal
catalyzes NO oxidation in
stripper.
1.0 (one absorption section)
Done prior to entrance to
absorber In a tower (unknown
dimensions)
None
None
1.3*
1.0
There appears to be some
indication that NO] oxidizes
carbon to CO or C02. Since this
would be accompanied by NO:
reduction, additional oxygen
Is required.
•v.20" H20
17,000
61
• MONSANTO RESEARCH CORPORATION •
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Monsanto process requires an acid-acid heat exchanger
(3^5,000 ft2) and 8 acid-water heat exchangers (total
of 45,000 ft ). Tyco process requires no acid heat
exchangers.
3. Detailed Comparison Using Identical Strippers
If a catalyst does exist which will effect the oxidation of
to N02 in the stripper, then the absorber is no longer required.
The process flow becomes much simpler. There is no need to have
acid-acid heat exchange or acid coolers, and the absorption tower
can be operated as a monofunctional tower. The dewatering and
humidifying sections can also be eliminated.
The Monsanto process and the Tyco.process thus become very similar.
The major differences are the following:
Monsanto oxidizes 862 in the absorber
Tyco oxidizes S02 in a separate chamber
Tyco process operates at 250°P requiring flue gas cooling
Monsanto process operates at . 300°F requiring flue gas
cooling
Tyco product acid concentration will be 80$
Monsantofs will be 87%.
P. Remaining Questions to be Answered
Will evaporative cooling eliminate the need to cool the
acid leaving the third section of the absorption tower?
What is the maximum gas flow rate that can be tolerated
in the absorption tower?
Will NO oxidation take place in the stripper and/or
absorber and eliminate the need for an oxidizing chamber?
What is the minimum acid rate required to cool and dewater
the flue gas?
Is it more economical to increase the size of the acid-acid
heat exchanger or to increase the temperature of the
stripping air?
Can flue gas be used as part of the stripping air (and
H20 requirement)?
Is it more economical to operate the stripper at high
temperature and low air rate or low temperature and.high
air rate?
62
• MONSANTO RESEARCH CORPORATION •
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What is the height of packing required in each of the
three sections of the absorption tower and stripping tower?
What will the pressure drop be through the absorption tower?
Is it necessary to monitor the NO/N02 ratio in the absorption
tower? If so, how?
What materials of construction are required?
Is there a catalyst that can be used in the stripper that
will permit oxidation of N203 to N02 and thereby facilitate
stripping of NOX at lower stripping temperatures?
63
• MONSANTO RESEARCH CORPORATION •
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REFERENCES
1. fieri, Z. Anorg. Allgem. Chem. 202, 113-34 (1931).
2. Tseitlin and Yavorskii, Journal of Applied Chemistry/U.S.S.R.
39 (5) (1966).
• MONSANTO RESEARCH CORPORATION •
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VI. PROCESS DESIGN
A. General
As a result of laboratory testing of catalysts of various types, a
process design evolved based upon commercial vanadia catalysts. It,
is necessary to note that the designs presented below were developed
solely from laboratory data supplemented with pertinent information
from the literature and various equipment vendors. There was no
piloting of the processes under the contract. It is also necessary
to note that the basic process.proposed here and labeled MRC/NAPCA
is superficially quite similar to published descriptions of Monsanto
Company's Cat-Ox** process. The two were developed completely
independently of each other. Published details of the Cat-Ox© process
only appeared toward the end of the present design effort. Although
such information was too late to be influential in the designs shown
below, it did serve very conveniently as a means of checking the
reasonableness of these designs at a number of points.
Four process designs are presented. The key step in each is
conversion of S02 to SC>3 over a supported vanadia catalyst at 850°F.
The basic process consists of such conversion followed by heat
recovery steps and then condensation of the sulfuric acid produced
for recovery. A minor departure from the basic derives from the
available temperature of flue gas. This normally is about 300°F.
for existing power boilers. The catalytic converter, however,
requires feed gas at 850°F. Consequently, the basic process applied
to an existing power plant requires adding provision for preheating
the gas fed to the converter.
In applying the process to a proposed power plant, provision can be
made in the power plant design to provide gas to the converter at
850° - 900°F. without the additional heating step.
A major departure in design, shown below, from the basic process
consists in the mode of S03 recovery proposed by the Gallery Chemical
Company.
The four cases of process designs presented, then, are intended to
provide a view of-the basic process, with various modifications,
under most probable applications.
Case I. MRC/NAPCA process applied to a large new
(1400 Mw) power plant (no separate preheat)
Case I-A. High temperature oxidation by solid catalysis
followed by reversible dry absorption of S03
(Gallery Chemical Company modification)
65
• MONSANTO RESEARCH CORPORATION •
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Case II. MRC/NAPCA process applied to a small (220 Mw)
existing power plant (separate preheat)
Case III. MRC/NAPCA process applied to the off-gas from
a reverberatory furnace of a small copper
smelter
Case I is illustrated in greater detail than the other three cases.
Capital and operating cost estimates are presented for all four
cases.
Cost estimates are based upon the following general assumptions for
all four cases. Special assumptions pertinent to a particular
process are presented along with the process descriptions below.
Flue gas and off-gas analysis:
Power Plant Smelter
Component. Flue-Gas (% by Volume) Off-Gas (% by Volume)
N2 74.9 76.75
C02 14.7 3.48
H20 7.25 8.94
'02 2.8 9.38
S02 0.3 1.45
NOX 0.05
Ply Ash 0.2 (by weight)
Following ratio factors were used for estimating the fixed
capital cost:
Piping 50$ of purchased equipment cost (except for
large new power plant facility)
Instruments 15$ of purchased equipment cost
Insulation 12$ of purchased equipment cost
Electrical 10$ of purchased equipment cost
Building 10$ of purchased equipment cost
66
• MONSANTO RESEARCH CORPORATION •
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Land and yard
improvements
Utilities
Engineering &
Construction
Contingency &
Contractor's
Pee
of purchased equipment cost
of purchased equipment cost
2055 of physical plant cost
20% of direct plant cost
Catalyst Life
Catalyst Cost
5 years
$l.l»5/liter
Amount of vanadia catalyst is based on W/F - conversion
profile for catalyst "A" at 850°F (Figure 32) plus 10$
excess catalyst
Size and shape of catalyst 3/8 in, cylinders
Rate of Return
Direct Labor
Supervision
Maintenance
Plant Supplies
Utilities
Payroll Burden
Plant Overhead
o% per year
$3.00/hr.
$7800 and $12,000 annually for first-
line supervisors and area super-
intendents, respectively
5% of the fixed capital investment
15% of the maintenance cost
a) steam — 50
-------
o
ui
H-
et
ut
Z
O
u
z
UJ
U
oc
IU
30
20
10-
CATALYST -A-
900°F
850°F
800°F
1.0 20
W/F, gm-sec/mole SO2><10'6
ao
Figure 32. W/F-Converslon Profiles for Catalyst 'A1
68
• MONSANTO RESEARCH CORPORATION •
-------
• Depreciation 10% of the fixed capital investment
• Taxes 2% of the fixed capital investment
• Insurance 1% of the fixed capital investment
Major equipment costs for a large new power plant facility (MRC/NAPCA
process) were substantiated by quotations from equipment vendors.
For other cases, equipment costs were estimated by using cost data
gathered by Monsanto's Central Engineering Department and cost data
from authoritative publications. Piping cost for the large new power
plant facility was estimated from labor and material take-off. For
this type of estimation, process flow sheet (Figure 33) and plan
and elevation drawings (Figures 31*, 35) were used.
Cost of equipment or other facilities normally required for usual
power plant operation is not incorporated in the capital investment
or operating cost estimates. Only the additional and/or incremental
costs for air pollution control are estimated.
B. Case I. MRC/NAPCA Process - New 1*100 Mw Plant
For a new power plant, the design assumes the availability of flue
gas at a temperature of 850° - 900°F. A flow diagram for one train
of this process is shown in Figure 33- There are four trains for
the 1*100 Mw station. Table 8 presents an overall material balance
and Table 9 shows the balance for a single train.
Hot gas from the boiler is relieved of 99+% of its dust burden in
a high temperature, electrostatic precipitator. The near perfect
efficiency of fly ash removal is necessary to minimize clogging, or
blinding, of the catalyst bed in the next step. The gas then flows
through fixed beds of vanadia catalyst where 90% of the sulfur
dioxide is oxidized to sulfur trioxide. The design of the
catalytic converter required optimizing pressure drop, through the
catalyst bed, with tower diameter. Tower diameter establishes gas
velocity which affects pressure drop. The latter relates directly
to operating cost. The converter design evolved consists of a set
of shallow catalyst beds in parallel. Gas flows up through the
beds, as indicated in Figure 35, to facilitate periodic cleaning
beds to remove fly ash. There are four such converters, each
with 13 beds for the gas handling capacity at this station size.
The program employed for converter optimization is given in Appendix
IVto this volume. Further, in Appendix III are the program and
results establishing the relationship between catalyst particle
geometry and pressure drop through the catalyst bed.
Gas leaving the converter is then cooled to about 500°F. In a new
power plant cooling is accomplished by stepwise passage through an
69
• MONSANTO RESEARCH CORPORATION •
-------
M
O
n
(A
o
o
•a
O
z
f'f. 1
CATALYTIC
CONVERTER
FROM
ELECTROSTATIC
PRECIPITATOR
200-225°F
TO
BOILER
COOLING WATER
WATER TO ECONOMIZER
FURNACE
PUMP ACID
TO
STORAGE
K-i
. ^-1 L.__
I
-- f-
*H-B
, ..'-.ir
OK«»-- I
H
MRC-fMPCA PROCESS
LARGE NEW POWER PLANT FACIL/Ty
1400 MW
SCAM
PROCESS FLOW DIAGRAM
OF
Figure 33. MRC-NAPCA Process, Large New Power Plant Facility, 1400 MW
Process Flow Diagram
-------
tf
13 Segments Each
9'-0' High
Catalytic Converter
Acid Tower & Mist
Eliminator
T-S' Dia. x OT-ff1
Acid Cooler
>. '** • '••I V
L' -NCFS
MAL«
I
—f~—
J««J
^..t*wArij,.«, r*T» •
MON»ANT<»
MRC-NAPCA PROCESS
LARGE NEW POWER PLANT FACILITY
I40OMW
PLOT PLAN
•ACILITY V. ..'
ri—
^^^^^1 T~Mkt
Figure 31*. MRC-NAPCA Process, Large New Power Plant Facility, 1*100 MW
Plot Plan
-------
I
3
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CatilyUc
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c
1 L-1
1 U
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MRC-NAPCA PROCESS
LARGE NEW POWER PL ANT FACILITY
1400 MW
ELEVATION
Figure 35. MRC-NAPCA Process, Large New Power Plant Facility, 1400 MW
Elevation Drawing
-------
Table 8
o
2
(/>
Z
O
m
(A
7)
O
I
O
O
a
TJ
O
SO; REMOVAL FROM FLUE GAS
CATEGORY: Large New Power Plant Facility
PROCESS FLOW SHEET
NAME OF PROCESS: MRC-NAPCA Process
MK:
FLUE GAS RATE: 2.5 x 106 SCFM
Process Stream
from boiler to converter.
Ib/hr
ll.SlxlO6
Flow
GPM SCFM
2.50X106
Temp.
°P
850-900
Composition
N2
70.27
CO;
21.66
H20
1.37
02
3.00
in Weight Percent
SO 2
0.61
NOx SO 3
0.05
HjSOfc CK<,
-— »
from converter to economizer 11.8lxlOe
from economizer to air heater ll.SlxlO6
from air heater to acid tower ll.SlxlO6
from acid tower to stack 11.67xl06
1.22 2.86 Trace 0.05 Trace 0.;^
4.22 2.86 Trace 0.05 Trace 0.?J
2.19xl06 865 70.27 21.66
2.t9xl06 665 70.27 21.66
2.')9xl06 150-500 70.27 21.66 1.22 2.86 Trace 0.05 Trace
2.17X106 200-225 71.07 21.91 3-98 2.89 Trace 0.05 Trace Trace
from acid tower to storage
13.26x10" 195.67
100
25.00
75-00
o
Z
-------
Table 9
SO; REMOVAL FROM FLUE CAS
o
CA
z
o
TO
5
PI
|
o
o
TO
3
CATEGORY: Large New Power Plant Facility
PROCESS PLOW SHEET
NAME OP PROCESS: MRC-NAPCA Process
Process Stream
from boiler to converter
from converter to economizer
from economizer to air heater
from air heater to "tee"
from "tee" to acid tower
from acid tower to stack
from acid tower to storage
Ib/hr
2.95x10*
2.95x10*
2.95x10*
2.95x10*
1.18x10*
1.146x10*
1. 66x10*
Plow
0PM SCPM
62.5x10"
62.3x10"
62.3x10"
62.3x10" •
31. 2x10"
30.9x10"
20.146
Temp.
op
850-900
865
665
150-500
1450-500
200-225
100
FLUE QAS RATE:
62.5 x
10" SCFM
Composition In Weight Percent
70.27
70.27
70.27
70.27
70.27
71.07
C02
21.66
21.66
21.66
21.66
21.66
21.91
H20
4.37
1.22
1.22
1.22
1.22
3-98
25.00
02
3.00
2.86
2.86
2.86
2.86
2.89
—
S02
0.61
Trace
Trace
Trace
Trace
Trace
NOx
0.05
0.05
0.05
0.05
0.05
0.05
—
SO 3
-__
Trace
Trace
Trace
Trace
Trace
H2SO* CH%
0.81
0.81
0.81
0.81
Trace
75.00
NOTE: Large stream of 2.5 x 10* SCPM Is divided Into four parallel streams
-------
economizer and air heater. The economizer takes heat from the flue
gas to heat boiler feed water while an air preheater is used to
recover more heat from the flue gas by heating the air prior to its
entry into the boiler firebox. Further cooling and condensation of
sulfuric acid mist is accomplished in packed acid towers which
operate in conjunction with acid coolers. Heat recovered in the
acid coolers is used to preheat boiler feed water. The acid towers
in this process, Figure 33, are reminiscent of the absorbers in a
contact sulfuric acid plant. In the latter, however, the primary
function is truly absorption, i.e., of S03 into slightly diluted
sulfuric acid where more sulfuric acid is formed. In the case of
the flue gas process, sulfuric acid.exists as such in the gas and
will condense as a mist if the gas is cooled. Consequently, the
main function of the acid towers here is to cool the flue gas to
the dew point of the acid which then condenses into the liquid in
the tower. The relationship of SOa in the flue gas to strength of
the acid condensing is ehowjTin Appendix, 10 .to this volume.
The fine sulfuric acid mist not retained in the acid tower is removed
by high efficiency (about 95%) mist eliminators prior to discharge
of the flue gas to the stack. Acid produced by this process is
stored for shipment as a 75 - 80$, or fertilizer grade, sulfuric
acid.
One of the major decisions in designing a process for hanging onto
the end of a power plant is the capacity for which to design. In
this case, we have designed for 100% capacity all of the time,
knowing, however, that this may not be realistic. On.the other
hand, designing for a lesser capacity would mean sacrificing emission
control during periods of peak capacity. In Figure 36 a scheme is
presented which permits the process to be run at incremental
capacities or at full capacity with only three of the four converters
on line. Each converter has a 10? overdesign. All four heat-
recovery trains are available, at all times to minimize loss of
thermal efficiency.
Figure 35 is an elevation view of one conversion train while Figure
34 is a plan view. Both figures give some idea of the size of the
treatment plant.
Instrumentation for the MRC/NAPCA process is comparatively simple
(Figure 37). Pressure drop through the catalytic converter indicates
fly ash accumulation in the catalyst beds. Instruments and control
equipment maintain proper temperature, pressure, draft and liquid
flow rates to maximize treatment and power plant efficiency and
minimize interference with the power plant operation.
Tables 10 through 14 present specifications for basic equipment
employed in this Case I process design.
75
• MONSANTO RESEARCH CORPORATION •
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CT\
-*- VALVES
1 Catalytic Converters
2 Economizers
3 Air Heaters
4 Acid Towers and Mist E liminators
5 I.D. Fans
6 Stacks
DECIMALS
Xk - I
xt.* i
»>I
-------
I
I
o
m
w
m
>
2
O
I
O
2
TJ
O
3
H
O
2
B
Flow Recorder
Flow Recorder Controller
Level Indicator
Pressure Indicator
Pressure Recorder
PDRC Press. Differential Recorder Controller
Tl Temperature Indicator
TR Temperature Recorder
Temperature Recorder Controller
FR
FRC
LI
PI
PR
TRC
— Process Piping Lead
Instrument Piping Lead
I nstru merit E lectrical Lead
->v- Instrument Air Lead
O Local Instrument
B Board-mounted Instrument
<8> Transmitter
£ Motor Control Valve
£ Hand Control Valve
TO
STACK
ELE
UNUSS OTHF.BWISf SPIOFICO
DIMENSIONS ARE I* INCHIS
TOLERANCES
DECIMALS FRACTIONS
»X = ± t
K XX - - ANGLES
XXXX BASIC i 3O
ALL SURFACES Vf
MATERIAL
FINISH
CMtCKtD
MON3ANTH ResE.AH4.il f'(
l>AVJ*fS I Mti'KAltm
1-A '. Itt.-H L IIIIO
MRC-NAPCA PROCESS
LARGE NEW POWER PLANT fKtUTY
I40OMW
SCALE «0(l«
INSTRUMENT FLOWSHEET
SHCl-T
B
Figure 37-
MRC-NAPCA Process, Large New Power Plant Facility, 1400 MW
Instrument Flowsheet
-------
Table 10
CATALYTIC CONVERTER
1. GENERAL
Number of reactors in parallel - 4
Number of catalyst beds in parallel in each reactor - 13
Overall height of each reactor - 117 ft.
Diameter of each reactor - 30 ft.
Height of each segment - 9 ft.
Height of each bed - 6-1/1 in.
Size and shape of catalyst - 3/8 in., cylinders
Total volume of catalyst - 20,000 cubic feet
2. OPERATING CONDITIONS OF EACH REACTOR
Total gas entering - 2.95 x 106 Ib/hr
Inlet temperature - 850°P
Outlet temperature - 865°P
Operating pressure - Slightly below atmospheric
Superficial gas velocity through each bed - 3-22 ft/sec.
Pressure drop - 0.96 in. H20
3. MATERIALS OF CONSTRUCTION
Converter shell - Carbon steel
Supports, beams, grates - Carbon steel (SA-285 Grade C)
Blank plates between each segment - Cast iron
78
• MONSANTO RESEARCH CORPORATION •
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Table 11
ACID RECOVERY AND MIST COLLECTION TOWER
1. GENERAL
Number of towers in parallel - 8
Overall height of each tower - 70 ft.
Diameter of each tower - 32 ft.
Packed height In each tower - 16 ft.
Size and type of packing - 3" Intalox Saddles
2. OPERATING CONDITIONS OF EACH TOWER
Total gas entering - 1.48 x 106 Ib/hr (31.2 x 10" SCFM)
Total liquid entering - 1.94 * 106 Ib/hr (2,320 GPM)
Gas inlet and outlet temperatures - 500°P and 225°F
Liquid inlet and outlet temperatures - 100°P and 225°F
Operating pressure - Slightly below atmospheric
Pressure drop through the acid recovery tower - 8 in.
Pressure drop through the mist eliminator - 10 in.
3. MATERIALS OF CONSTRUCTION
Tower - Carbon-steel vessel with minimum of 4" of acid-proof
brick lining.
Packing - Chemical stoneware
Mist element - Chemically resistant glass fibers
79
• MONSANTO RESEARCH CORPORATION •
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Table 12
ACID COOLER
1. GENERAL
Size - 7' - 6" 0 x 20'
Type - Shell and Tube
Sq. Ft. Surface/Shell - 27,750
Number of Units - 8
2. PERFORMANCE OF ONE UNIT
Fluid Circulated
Total Fluid Entering
Temperature In
Temperature Out
Pressure Drop
Shell Side
Water
1.125 x 106 Ib/hr
(2,250 GPM)
80°F
181°F
20 PSI
Tube Side
75% Sulfurlc Acid
1.9^ x 106 Ib/hr
(2,320 QPM)
100°F
225°F
10 PSI
Heat Exchanged - 114 x 106 Btu/Hr
MTD (Corrected) - 25.2°F
Transfer Rate Service - 163 Btu/(Hr) (ft2)(°F)
3. MATERIALS OF CONSTRUCTION
Tubes: Welded Carpenter 20
Shell: Carbon Steel
80
• MONSANTO RESEARCH CORPORATION •
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Table 13
ACID PUMP
1. GENERAL
Pump Type - Centrifugal
Number of Units - 8
2. OPERATING CONDITIONS OF EACH UNIT
Liquid Pumped: 75% Sulfuric Acid
Capacity, % 225°F, Normal 2,300 GPM, Design 2,540 0PM
Specific Gravity % 225°F - 1.6?
Viscosity § 225°P - 6 cps
Diff. Head - 65 ft.
3. MATERIALS OF CONSTRUCTION
Casing - Stainless Steel, Alloy 20
Impeller - Stainless Steel, Alloy 20
Shaft - Stainless Steel, Alloy 20
Packing - None
4. DRIVER DATA
Motor HP - 100, RPM - 1160
Phase - 3, Cycle - 60, Volts - 440
81
• MONSANTO RESEARCH CORPORATION •
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Table 14
INDUCED DRAFT PAN
1. GENERAL
Type of Fan - Double Inlet Centrifugal Fan
Number of Units - 8
Size: Height - 13 ft, Width - 18 ft, Length - 10 ft
2. OPERATING CONDITIONS OF EACH UNIT
Inlet Temperature - 225°F
Suction Pressure - 30 In. Water
Qas Density - 0.06 lb/ft3
Capacity - 400,000 CFM % 225°F and 14.7 psla
3. TYPE OF FLOW CONTROL
Discharge Damper
4. MATERIAL OF CONSTRUCTION
Carbon Steel
5. DRIVER DATA
Motor HP - 3,000, RPM - 1,200
Phase - 3, Cycle - 60, Volts - 2,300
Accessory - Brushless type exciter
82
• MONSANTO RESEARCH CORPORATION •
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In addition to the general assumptions previously cited, the
following assumptions were also employed in the Case I capital
and operating cost estimates:
Flue Gas Rate
Size of Power Plant
Coal Required for
Power Plant
Operating Factor
S02 Conversion
Recovered as
Flue Gas Temperature
Cost of 3" Intalox
Saddles
2.5 MM SCFM
1400 Mw
580 tons/hr
330 days/year § 100$ capacity
90%
95%
850 to 900°F
$4.55/ft3
Tables 15 through 18 show capital and operating cost estimates: for
Case I. Figure 38 shows the effect of product sale price on net
operating cost.
C. Case I-A. Reversible Dry Absorption of SOg
This case differs from Case I primarily in the mode of product
recovery. In this process, sulfur dioxide in flue gas is first
oxidized to sulfur trioxide over 'vanadia catalyst, as in Case I.
However, the product sulfur trioxide is removed from the main gas
stream by sorption on a dry medium, Na^Oi^, at relatively high
temperature. Desorption is effected at still higher temperature
by decomposition of the sorption product. The reactions involved
are:
S03
850°F
Na2S207
1000°F
Na2S207 = Na2SOi« + S03
83
• MONSANTO RESEARCH CORPORATION •
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30
£
3
**>
i
i
8.
4*
810
Q.6
Break Even Point
Mills per Kilowatt Hour
Ol2 ! 0.2
1.0
03 0 Oi5
$ per Ton of Coal
1.0
15
Figure 38. Effect of Product Credit on Operating Cost of
MRC/NAPCA Process (Large New Power Plant Facility)
• MONSANTO RESEARCH CORPORATION •
-------
Table 15
CAPITAL COST ESTIMATE SUMMARY
Category: NEW POWER PLANT (HIGH TEMPERATURE EFFLUENT)
Case HI
Name of Process: MRC-NAPCA Process Flue Gas Rate; 2.5 MMSCFM
1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
21.
MW 1,100
Purchased Equipment
Purchased Equipment Installation
Piping
Instruments
Insulation
Electrical
Building
Land & Yard Improvements
Utilities
Incremental Cost of Precipitator
Acid Tower Packing
Mist Elements
Initial Catalyst Cost
Physical-Plant Cost
Engineering and Construction
Direct Plant Cost
Contingency & Contractor's Fee
Fixed Capital Cost
Capitalized Cost of Catalyst
Working Capital
TOTAL INVESTMENT
Capital Requirements
$/KW Capacity 21-71
Cost - $
6,632,000
2,539,000
2,653,000
995,000
800,000
663,000
663,000
995,000
995,000
1,500,000
155,000
2,3^0,000
630,000
21,860,000
1,372,000
26,232,000
5,216,100
31,178,100
2,110,000
1,021,000
31,639,100
% of Total
19.15
7.33
7.66
2.87
2.30
1.91
1.91
2.87
2.87
1.31
1.31
6.76
1.82
63.10
12.62
75.72
15-15
90.87
6.18
2.95
100.00
85
• MONSANTO RESEARCH CORPORATION •
-------
Table 16
EQUIPMENT COST ESTIMATE SUMMARY
Category: NEW POWER PLANT (HIGH TEMPERATURE EFFLUENT)
• ' •
2 Name of Process: HRC-NAPCA Process Flue Gas Rate: 2.5 MMSCPM
o
MU 1*00
m
CD
m
m Item _ No. of Units Cost -
- - —
5
o
I
o
o
a
•o
o
a
>
H
O
1.
2.
3.
H.
5.
6.
Catalytic Converters 4
Acid Towers & Mist
Acid Coolers
Acid Pumps
I.D. Fans
Storage Tanks
Eliminators
8
8
8
8
2
792,000
2,080,000
2,780,000
64,000
720,000
196,000
Purchased Equipment Cost 6,632,000
-------
Table 17
WORKING CAPITAL ESTIMATE SUMMARY
o
z
(/>
z
o
V
m
(/>
m
>
TO
O
I
O
o
33
TJ
O
a
>
H
O
CO
-J
Category: NEW POWER PLANT (HIGH
TEMPERATURE EFFLUENT)
Name of Process :MRC-NAPCA Process Flue Gas Rate: 2.5 MMSCFM
Item
1.
2.
3.
4.
5-
6.
7.
MW: 1400
Raw material inventory, 1 Month
Direct labor, 3 Months
Indirect cost, 3 Months
Operating supplies, 3 Months
Fixed costs
Spare parts
Miscellaneous
Cost - $
13,200
217,000
235,600
55,600
222,300
185,300
92,000
Percent
1.30
21.25
23.08
5.44
21.78
18.14
9.01
TOTAL
1,021,000
100.00
-------
Table 18
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year 6 100JE Capacity
Category: NEW POWER PLANT (HIGH TEMPERATURE EFFLUENT)
Name of Process MRC-NAPCA ProcessFlue Gas Rate
2.5
MMSCFM
1.
2.
3-
1.
5.
6.
7.
8.
9.
10.
11.
12.
13-
11.
15.
16.
17.
18.
19.
20.
21.
22.
23.
21.
MW 1,100
Fixed Capital Cost
ITEM
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,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed
Capital/Yr.
Taxes, 255 of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 1-78
Mill/kwh 0.7371
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mlll/kwh
31,178,100
TOTAL $
219,000
105,000
23,^00
1,573,900
236,090
895,700
_
3,083,090
25,700
969,200
_
_
_
991,900
3,117,810
62Q.S70
311,780
_
i.oqp.iqo
8,170,180
PER CENT
3-05
1.28
0.28
19-27
2.89
10.97
-
37.71
0.31
11.86
-
-
-
12.17
38.53
7.71
3-85
-
50.09
100.00
• MONSANTO RESEARCH CORPORATION •
-------
The desorbed sulfur trioxide is fed to an absorption tower
essentially the same as that in a contact sulfuric acid plant,
to make 100$ sulfuric acid and oleum.
As shown in the flow diagram in Figure 39, flue gas at high tempera-
ture (850° to 900°F) flows through a catalytic converter where 90$
of the sulfur dioxide is converted to sulfur trioxide. Next it
passes through an absorber-stripper where S03 is absorbed on a dry
medium consisting of porous silica or alumina granules impregnated
with 20% sodium sulfate. Sulfur trioxide is stripped from exhausted
sorbent with hot gas from a furnace at 1000°F. The absorber-
stripper is a fixed bed contactor designed very much like the
converter to minimize pressure drop through the absorbent beds.
Assuming the absorbent requirement is the stoichiometric amount,
each absorber-stripper contains 13 bed sections in parallel and is
30 ft. in diameter by 117 ft. in height.
Four absorption towers are required to permit adequate cycle, timing.
This is illustrated by the following.table:
ABSORBER NUMBER
15 min abs.
15 min des. 15 min abs.
15 min des. 15 min des. 15 min abs.
15 min cool 15 min des. 15 min des. 15 min abs
Clean, hot gases emerging during the absorption cycle, at.about
820°F, are passed through an economizer and an air heater where
heat is recovered for utilization in the power plant. An induced
draft fan moves the clean gas through the system and up the stack.
The hot gas emerging during the desorption cycle is cooled to about
500°F in a gas cooler. Lower temperature would result In premature
condensation of acid. The cooled gas is then directed to an acid
recovery unit for producing concentrated acid. The sulfur trioxide
concentration in this gas stream is high, as shown by the material
balance in Table 19- Thus, the primary function of the acid tower^,
in this case, is absorption of S03 into 96-98$ sulfuric acid. Mist
eliminators are, again, required on the tail gas from the absorber.
The interesting point is that the gas rate to the absorber is two
orders of magnitude less than the main gas stream flow. Consequently,
only a single acid tower is required compared to eight in Case I.
89
• MONSANTO RESEARCH CORPORATION •
-------
O
z
Ul
z
O
TO
CJ
ra
o
o
x
3
a
H
O
Absorber-Stripper
820>F
To Hot To
Boiler Water Boiler Air
620°F
300°F
Catalytic
Converter
From
Electrostatic
Precipitator
LQ
• A.
r 1
'.»'*», -*r: •wn^.ii.
i.* ¥tNSf«'Ni «•'? .'« ',''•'.*)•?
rc;.m»N« i'-"
oec'M'«ts H*t:Tu>t«s
»x
xxx> B*^' : . « jo
• I I S'JRf«Cf « X''
MA-»(_-<«!A..
«• •"i-il.
(MK-n
Al»»0
»**Fj
AP^fJ
f.-*ri«o
Utt AV. '4
f
—t-
(
U. -i .-
i
MOVHA.Vr«.> Hi>-T. >i»t 1! t ••«!•«. HAT ION
-.H1..X , *, i: , •... . \
us \ t i • V •'
Reversible Dry Absorbent Process ! prorjK^ Fk» Dianram
Laroe New Power Plant Facility Process now oiagrai
1400 MW >
•••-•-• >-Nonej-v ;;, ( : i ^^j [V "CT"
K-E -.-:.-.:
Figure 39. Reversible Dry Absorbent Process, Process Plow Diagram
-------
Table 19
SO, REMOVAL FROM FLUE GAS
CATEGORY: Large Hew Power Plant Facility
PROCESS FLOW SHEET
NAME OF PROCESS: Reversible Dry Absorbent
MW: HOO
FLUE GAS RATE: 2.5 x 106 SCFK
O
•z.
TL
H
O
a
m
u>
m
o
i
o
o
TO
•o
O
TO
O
Z
Process Stream
from electrostatic precipltator to
catalytic converter
from catalytic converter to absorber-
stripper
from absorber-stripper to economizer
from furnace to absorber-stripper
from absorber-stripper to gas cooler
from gas cooler to acid tower and
mist eliminator
from acid tower and mist eliminator
to I.D. fan
from economizer to air heater
from air heater to I.D. fan
from I.D. fan to stack
from acid tower to storage
Ib/hr
ll.SlxlO6
ll.SlxlO6
11.73xl06
O.llSxlO6
0.199xl06
0.199xl06
O.llSxlO6
11.73xl06
11.73xl06
11.817xl06
n i *i n&
Flow
GPM
2
2
2
2
3
3
2
2
2
2
i n? ii
SCFM
.50xl06
•50xl06
.igxio6
.5x10"
. 1x10"
.1x10"
.5x10"
.igxiu6
.I9xl06
•5xl06
Temp.
°F
850-900
865
820
1000
1000
150-500
225
620
300
300
i nr,
Composition
N,
70.27
70.18
70.69
70.18
11.60
11.60
70.19
70.69
70.69
70.69
C02
21.66
21.61
21.79
21.61
12.82
12.82
21.61
21.79
21.79
21.79
H2
1.
14.
1.
1.
2.
2.
1.
1.
1.
1.
n
0
37
10
15
10
61
61
10
15
15
15
t;
°2
3.00
2.93
2.96
3.07
1.82
1.82
3.08
2.96
2.96
2.95
in Weight Percent
S02
0.61
0.06
0.06
0.65
0.38
0.38
0.61
0.06
0.06
0.07
0.05
0.06
0.06
0.06
0.03
0.03
0.05
0.06
0.06
0.05
S03 KjSO,,
0.73 —
Trace
10.71
10.71
Trace
Trace
Trace ..
Trace
QQ ^
-------
This difference is reflected in the comparative economics of cases
I and I-A. Case I- A still requires four converters, but only one
acid tower.
Additional assumptions employed in the cost estimates for Case I-A
are:
Absorption Efficiency a) 100$, i.e., stoichiometric,
for 15 minutes
b) 50% stoichiometric for 15
minutes
Desorption Efficiency 100$ in 30 minutes
• Amount of NazSOi* in 20$ (by weight)
total absorbent
Cost of Absorbent 5 shows the effect of product credit on net
operating cost.
92
• MONSANTO RESEARCH CORPORATION •
-------
Table 20
CAPITAL COST ESTIMATE SUMMARY
Name
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
21.
Category: NEW POWER PLANT (HIGH
Case #1A
of Process: Reversible Dry Absorbent
MW moo
Purchased Equipment
Purchased Equipment Installation
Piping
Instruments
Insulation
Electrical
Building
Land & Yard Improvements
Utilities
Incremental Cost of Precipitator
Acid Tower Packing
Mist Elements
Initial Catalyst Cost
Physical-Plant Cost
Engineering and Construction
Direct Plant Cost
Contingency & Contractor's Pee
Fixed Capital Cost
Capitalized Cost of Catalyst
Working Capital
TOTAL INVESTMENT
Capital Requirements
$/KW Capacity 12-^8
TEMPERATURE EFFLUENT)
Flue Gas Rate: 2.5
Cost - $
2,934,000
1,400,000
1,470,000
440,000
353,000
294,000
294,000
440,000
440,000
1,500,000
30,000
"668,000
10,263,000
2,053,000
12,316,000
2,463,000
14,779,000
2,140,000
551,000
17,470,000
MMSCFM
% of Total
16.79
8.01
8.41
2.52
2.02
1.68
1.68
2.52
2.52
8.59
0.17
—
3.82
58.75
11.75
70.50
14.10 .
84.60
12.25
3,15
100.00
93
• MONSANTO RESEARCH CORPORATION •
-------
Table 21
EQUIPMENT COST ESTIMATE SUMMARY
Category: NEW POWER PLANT (HIGH TEMPERATURE EFFLUENT)
Name of Process: Reversible Dry Absorbent Flue Gas Rate: 2.5 MMSCFM
MW 1400
•
o
z
u»
z
o Item Cost - $
a
w 1. Catalytic Converters 792,000
£ vo 2. Absorber-Strippers 792,000
o •*=•
1 3- Furnace 200,000
§ 4. Gas Cooler 65,000
o 5. Absorber and Mist Eliminator 205,000
H 6. Acid Cooler 50,000
z 7. Regenerator Blower 70,000
* 8. I.D. Fans 610,000
9. Storage Tanks 150,000
PURCHASED EQUIPMENT. COST 2,93^,000
-------
Table 22
WORKING CAPITAL ESTIMATE SUMMARY
o
z
z
H
O
3)
m
M
m
> vo
3D U1
O
Z
o
o
a
TJ
o
O
Z
ame
Category: NEW POWER PLANT (HIGH TEMPERATURE
of Process: Reversible Dry Absorbent Flue
MW: moo
Item
1.
2.
3-
4.
5.
6.
7.
Raw Materials & Chemicals Inventory f 1 Month
Direct Labor, 3 Months
Indirect Cost, 3 Months
Operating Supplies, 3 Months
Fixed Costs
Spare Parts
Miscellaneous
TOTAL
EFFLUENT)
Gas Rate: 2.5
Cost, $
70,000
113,000
115,600
24,300
97,100
81,000
50,000
551,000
MMSCFM
Percent
12.71
20.51
20,98
4.41
17.62
14.70
9.07
100.00
-------
Table 23
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year g 100$ Capacity
Category: NEW POWER PLANT (HIGH TEMPERATURE EFFLUENT)
Name of Process: Reversible Dry Absorbent Flue Gas Rate: 2.5 MMSCFM
MW 1100
Fixed Capital Cost: $11.779.000
1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13-
15.
16.
17.
18.
19.
20.
21.
22.
23-
21.
ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15? of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20$ of 2 & 3
Plant Overhead, 50$ of 2, 3,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10$ Fixed
CapHaT/Yr .
Taxes, 2$ of Fixed Capital
Insurance, 1$ of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 1.013
Mill/kwh U.132
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh
TOTAL $
731,000
me; nnn
23,100
739,000
110,800
615,000
2,351,200
25,700
189,100
_ —
511,800
M77,900
295, 6$0
117, 8QO
1,921,300
1,790,3o'o
PER CENT
15.26
2.1Q
0.19
15.13
2.31
13.16
___
19.15
0.51
10.21
— __
— — _
___
10.75
30.85
6.17
3-08
10.11
100.00
96
• MONSANTO RESEARCH CORPORATION •
-------
40-i
30-
to
*h_
3
20-
el-
's
Break Even Point
8 10-
0
0.6
0.4
Mills per Kilowatt Hour
0.2
0.2
0.6
1.5
1.0
0.5
$ per Ton of Coal
0.5
1.0
Figure 40. Effect of Product Credit on Operating
Cost of Reversible Dry Absorbent Process
(Large New Power Plant Facility)
97
• MONSANTO RESEARCH CORPORATION •
-------
«
§
I
5
i
3
co
TO STACK
FROM
ELECTROSTATIC
PRECIPITATOR
ACID TOWER
AND MIST
ELIMINATOR
474°F
100°F
ACID
ER
I.D. FAN
• SULFUR 1C ACID TO STORAGE
WATER IN
WATER OUT
ACID
PUMP
COMBUSTION
GAS
•jMi :* OTHfo«fl«f Vt^.r.ri;
p'vt -.^' -. , »wf is (Nt. <«r%
TO_H»M».-f »
ur"" V*LV. r-(»f:<-v*
• V
» 1 X ' B*«. ' . • JO
Al.l. »U •:'»*£* .
MATtB.A. ....
«»»C |
»«
AVf-.. 1
»»o ~!
..^•.^T
- t ...&!•. f- -
t~ •— -
::•"»
MOXS-XNTO Kji.^rAIU tl
1 , 'X 1 M«t it »
. •. ' X. • I . >
MRC-NAPC A PROCESS
4MAI 1 FYRTMfi PtlU/FP PI AUT MTI ITV
220 MW
«.:.-,Fliae:v, -'.-
1 ~ U__
( '• :!• I*«IIA I'|I>.N
PROCESS FUMf OM8RMM
»•. •>-. j«i»
S'ifF.I ;tf
Figure 4l. MRC-NAPCA Process, Small Existing Power Plant Facility, 220 MW
Process Flow Diagram
-------
Table
30? REMOVAL FROM FLUE HAT-
2
o
z
.z
H
O
K
m
o
x
0
o
a
•o
o
vo
CATEGORY:
NAME OP PROCESS: MRC-NAPCA Process
Process Stream
from boiler to flue gas heat exchanger
from flue gas heat exchanger to burner
from burner to catalytic converter
from catalytic converter to flue gas heat exchanger
from flue gas heat exchanger to acid tower
from acid tower to stack
combustion gas to burner
product acid to storage
2.
2.
2.
2.
2.
2.
0.
26
Ib/hr
36x10'
36x10'
58x10'
58x10'
58x10'
56x10'
22x10'
.5xl03
Snail Existing Power Plant Facility
PROCESS FLOW SHEET
HW: 220 FLUE GAS RATE: 0.5
Plow
GPM
0.
0.
0.
0.
0.
0.
0.
39.13
SCFM
50x10'
50x10'
55x10'
55x10'
55x10'
55x10'
05x10'
Temp.
op
300
683
850
865
171
225
100
x 10' SCFM
Composition in Weight Percent
N2
70
70
70
70
70
71
73
CO 2
.27
.27
.52
.52
.52
.25
.20
21
21
20
20
20
21
_
.66
.66
.87
.87
.87
.09
H2O
1.37
1.37
1.69
1.71
1.71
1.53
25.00
O2 S02
3
3
3
2
2
2
_
.00
.00
.08
.95
.95
-98
22
0.
0.
0.
0.
0.
61
61
59
06
06
Trace
_
N'0X SO 3
0.05
0.05
0.05
0.05 0.03
0.05 0.03
O.on Trace
—
h2SOi, CKw
___
0.77
0.77
I
Trace
75.00
O
z
-------
Table 25
CAPITAL COST ESTIMATE SUMMARY
Category: SMALL EXISTING POWER PLANT
Case #2
Name of Process: MRC-NAPCA Process Flue Gas Rate: 0.5 MMSCFM
MW 220
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
21.
Purchased Equipment
Purchased Equipment Installation
Piping
Instruments
Insulation
Electrical
Building
Land & Yard Improvements
Utilities
Incremental Cost of Precipltator
Acid Tower Packing
Mist Elements
Initial Catalyst Cost
Physical-Plant Cost
Engineering and Construction
Direct Plant Cost
Contingency & Contractor's Fee
Fixed Capital Cost
Capitalized Cost of Catalyst
Working Capital
TOTAL INVESTMENT
Capital Requirements
$/KW Capacity 16.12
Cost - $
2,011,000
770,500
1,007,000
302,100
211,700
201,100
201,100
302,100
302,100
500,000
51,000
520,000
126,100
6,512,700
1,308,600
7,851,300
1,570,300
9,121,600
126,000
365,200
in ?i? ftnn
% of Total
19.73
7.55
9.87
2.95
2.37
1.98
.1.98
2.95
2.95
1.89
0.52
5.09
1.23
61.06
12.81
76.87
15.38
92.25
1.18
3.57
100 00
100
• MONSANTO RESEARCH CORPORATION
-------
o
z
H
m
CO
Table 26
EQUIPMENT COST ESTIMATE SUMMARY
Category: SMALL EXISTING POWER PLANT
Name of Process: MRC-NAPCA Process Flue Gas Rate: 0.5 MMSCFM
MW 220
-
Item Cost - $
a 2 Catalytic Converter 278,000
o *~^
1 Acid Tower & Mist Eliminator 340,000
° Acid Cooler 540,000
o Acid Pump 12,500
H Flue Gas Heat Exchanger 595,000
z I. D. Fan 180,000
* Flue Gas Burner 8,500
Storage Tank 60,000
Purchased Equipment Cost 2,014,000
-------
Table 27
WORKING CAPITAL ESTIMATE SUMMARY
Category: SMALL EXISTING POWER PLANT
Name of Process: MRC-NAPCA Process Flue Gas Rate: 0.5 MMSCFM
* MW 220
z
w
z
-I
o
a Item Cost - $ Percent
m
> o
21 ro
O
X
o
o
3)
a
H
O
•z
1.
2.
3.
"•
5.
6.
7.
Raw Material Inventory
1 Month
Direct Labor, 3 Months
Indirect Cost, 3 Months
Operating Supplies
3 Months
Fixed Costs
Spare Parts
Miscellaneous
36,200
78,600
79,700
16,500
66,000
55,000
33,200
9.92
21.52
21.82
4.51
18.07
15.06
9.10
Total 365,200 100.00
-------
Table 28
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year € 60jT Capacity
Category: SMALL EXISTING POWER PLANT
Name of Process MRC-NAPCA Process Flue Gas Rate
0.5
MMSCFM
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15-
16.
17.
18.
19.
20.
21.
22.
23-
24.
MW 220
Fixed Capital Cost 9
ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 151 of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20$ of 2 & 3
Plant Overhead, 50? of 2, 3,
4 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, IQ % Fixed
Capital/Yr.
Taxes, 2% of Fixed Capital
Insurance, 1$ of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 6.51
Mill/kwh 2.66
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh
,421,600
TOTAL $
444,250
79,000
15,600
471,080
70,700
141,000
.-
1,221,630
18,920
318,190
-
-
-
337,110
942,160
188,440
94,220
_
1,224,820
?)7fl^,'i60
PER CENT
15.96
2.84
0.56
16.93
2.53
5.07
-
43.89
0.67
11.44
-
-
-
12.11
33.85
6.77
3-38
-
44.00
100.00
103
• MONSANTO RESEARCH CORPORATION •
-------
60i
50-
40-
30
Q_
"o
C
o
8
Profit
0.5
Break Even Point
Cost
Mills per Kilowatt Hour
0,5 1,0 1,5
3.0
To
1.0 2.0 3.0
$ per Ton of Coal
4.0 5.0 6.0 7.0
Figure 42. Effect of Product Credit on Operating Cost of
MRC^NAPCA Process (Small Existing Power Plant
Facility)
• MONSANTO RESEARCH CORPORATION •
-------
Additional assumptions employed in material balance and cost.
estimates for Case II are:
. Flue Gas Rate 0.5 MM SCPM
• Size of Power Plant 220 Mw
• Coal Required for 90 tons/hr
Power Plant
• Operating Factor 330 days/year @ 60% capacity
• SO 2 Conversion 90%
• SO 3 Recovered as 95%
• Flue Gas Temperature 300°F
• Cost of 3" Intalox $I».55/ft3
Saddles
• Cost of Natural Gas 35
-------
in height, and one acid tower, 20 ft. in diameter by 46 ft. in
height. The converter consists of nine catalyst beds in parallel.
For this case, the pertinent assumptions, over and above the general
ones cited earlier, are as follows:
Off-Gas Rate
Copper Production
Capacity
Operating Factor
SOa Conversion
303 Recovered as
Off-Gas Temperature
Cost of 3" Intalox
Saddles
90,800 SCFM
230 tons/day
330 days/year § 100? capacity
90$
95$
$4.55/ft3
Here, as in all previous cases, the amount of vanadia catalyst is
derived from Figure 32. Tables 30 through 33 show capital and
operating cost estimates and Figure 4*1 shows the effect of product
sale price on net operating cost.
106
• MONSANTO RESEARCH CORPORATION •
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S 3 ,
B
Acid Tower
and Mist
Eliminator
Offgas
Electrostatic Heat
Precipitator Exchanger
Offgases
From
Smelting
Plant
Add
Cooler
450°F
850°F
REVISIONS
'•W
Sulfuric Acid to Storage
Water In
Water Out
Catalytic
Convwter
TOLCB»KCCS
KX -- J* '
xxx -- * ANUI t:
xr »* BA?'C - »»o'
•ii si^r»cE» i/
m»rc» »L
MtlNSANfl-
r i > . -• :. »IM tti v :\.»i v
MRC-NAPCA PROCESS
SMELTER BJCILTTY
PROCESS FU
Figure 43. MRC-NAPCA Process, Smelter Facility
-------
Table 29
Z
o
m
(A
5
x
O
X
o
o
a
•o
o
5
CD
S02 REMOVAL FROM FLUE CAS
CATEGORY: Smelter Facility
PROCESS PLOW SHEET
NAME OF PROCESS: MRC-NAPCA Process
Flow
Process Stream
from smelting plant to electrostatic
•* precipltator
from electrostatic precipltator to
heat exchanger
from heat exchanger to converter
from converter to heat exchanger
from heat exchanger to acid tower
from acid tower to stack
product acid to storage
Ib/hr
0.13x10*
0.13x10*
0.13x10*
0.13x10*
0.13x10*
0.11x10*
21,600
0PM SCFM
90
90
90
89
89
85
36.3 -
,800
,800
,800
,081
,081
,913
—
Temp.
op
150
150
850
918
523
218
100
Composition
N2
75.25
75.25
75.25
75.25
75.25
79.77
C02
5-36
5.36
5-36
5-36
5.36
5.68
H20
5.61
5.61
5.61
1.85
«.85
3.61
25.00
02
10.50
10.50
10.50
9.77
9-77
10.36
FLUE
GAS RATE: 90,800
SCFM
in Weight Percent
SO 2
3.25
3.25
3.25
Trace
Trace
Trace
NOx SO 3 M2SO»
Trace 1.25
Trace 1.25
— Trace Trace
75.00
CH»
-------
Table 30
CAPITAL COST ESTIMATE SUMMARY
Category: SMELTER FACILITY.
Case # 3
Name of Process: MRC-NAPCA Process Off-gas Rate; 90.800 SCFM
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
•13.
11.
15.
16.
17.
18.
19.
20.
Purchased Equipment
Purchased Equipment Installation
Piping
Instruments
Insulation
Electrical
Building
Land & Yard Improvements
Utilities
Incremental Cost of Precipitator
Acid Tower Packing
Mist Elements
Initial Catalyst Cost
Physical-Plant Cost
Engineering and Construction
Direct Plant Cost
Contingency & Contractor's Pee
Fixed Capital Cost
Capitalized Cost of Catalyst
Working Capital
TOTAL INVESTMENT
Cost - $
1,306,000
524,700
653,000'
195,900
156,800
130,6.00
130,600
195,900
195,900
'
12,000
226,000
164,000
3, 891, 400
778,300
4,669,700
934,000
5,603,700
553,500
214,730
6,371,930
% of Total
20.50
8.24
. 10.25
3.08
2.46
2.04
2.04
3.08
3.08
__
0.18
3-55
2.57
61.07
12.21
73.28
14.66
87.94
8.69
3.37
100.00
109
• MONSANTO RESEARCH CORPORATION •
-------
o
z
M
Table 31
EQUIPMENT COST ESTIMATE SUMMARY
Category: SMELTER FACILITY
Name of Process: MRC-NAPCA Process Flue Gas Rate: 90.800 SCFM
a
m
at
> £ Item Cost -$
n o
I
o Electrostatic Precipitator ^30,000
o
a)
3 Catalytic Converter 1^3,000
a
H Acid Tower & Mist Eliminator 320,000
o
1 Acid Cooler '96,000
Acid Pump 2,000
I.D. Fan ^5,000
Off-gas Heat Exchanger 205,000
Storage Tank 65,000
Purchased Equipment Cost 1,306,000
-------
o
z
z
H
O
70
m
u>
m
>
a
o
i
o
o
O
OJ
O
Z
Table 32
WORKING CAPITAL ESTIMATE SUMMARY
Category: SMELTER FACILITY
Name of Process: MRC-NAPCA Process Off Gas Rate: 90,800 SCFM
Item
1.
2.
3.
4.
5.
6.
7.
Raw Material Inventory, 1 Month
Direct Labor, 3 Months
Indirect Cost, 3 Months
Operating Supplies, 3 Months
Fixed Costs
Spare Parts
Miscellaneous
Cost - $
3
52
52
10
12
35
19
,420
,050
,200
,500 .
,030
,030
,500
Percent
1
24
. 24
4
19
16
9
.60
.24
.30
.88
.58
.32
.08
TOTAL 214.730 100.00
-------
Table 33
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year g 100$ .Capacity
Category: SMELTER FACILITY
Name of Process MRC-NAPCA Process Off-ga§ Rate 90,800 SCFM
1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
Fixed Capital Cost
ITEM
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,
1 and 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
5,603,700
TOTAL $
in ,onn
R?,^no
1R.600
280,200
1?,000
18,^00
__
1lQr600
-,-, finn
195.200
__
__
— —
208.800
560,100
112,100
56,000
—
728,500
1.386.900
PER CENT
2.96
3.79
1.12
20.21
3.03
1.31
— _
_^2.A2_
0.98
11.07
—
—
—
15.05
10.11
8.09
1.03
—
52.53
100.00
112
• MONSANTO RESEARCH CORPORATION •
-------
0.
15
Break Even Point
U) 0.5 0 03
Net Operating Cost, 10*Dollars/Year
1.0
1.5
Figure
Effect of Product Credit on
Operating Cost of MRC-NAPCA
Process (Smelter Facility)
113
• MONSANTO RESEARCH CORPORATION
-------
REFERENCES
1. Arthur G. McKee and Company, Systems Study for Control of
Emissions - Primary Nonferrous Smelting Industry, Vol. Ill
(Final Report), Contract PH 86-65-85, Plate No. C-l,
San Francisco, June 1969.
• MONSANTO RESEARCH CORPORATION •
-------
APPENDIX I
COMMERCIAL CATALYST TEST DATA
115
• MONSANTO RESEARCH CORPORATION •
-------
100
CATALYST -A-
10 2:0
W/F, gin-sec/mole SO2><10~6
Figure 45. W/F-Conversion Profile for Catalyst
'A' at 900°F
116
,• MONSANTO RESEARCH CORPORATION •
-------
100
10
CATALYST -A-
1.0 2.0
W/F, gm.»ec/mole SO2><10~6
3.0
Figure 46. W/F-Conversion Profile for Catalyst
'A' at 850°F
117
• MONSANTO RESEARCH CORPORATION •
-------
100i
90-
80-
o
iu
70
z
o
u
o
m
Z
iu
U
5(
Catalyst A
40-
30-
20-
10-
1.0 2.0
W/F, gm-sec/mole SO2xlO~6
3.0
Figure ^1. W/P-Conversion Profile for Catalyst
'A' at 800°F
118
• MONSANTO RESEARCH CORPORATION •
-------
Table 3^
CATALYST -'ATTEST RESULTS
Run
No.
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
* S02
in feed
(mole )
0.34
0.34
0.31
0.30
0.34
0.319
0.346
0.280
0.281
0.296
0.300
0.292
0.318
0.282
0.276
0.308
0.302
0.291
0.284
0.322
Catalyst
Weight
(g)
6.0
6.0.
6.0
6.0
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
Total
Flow
(cc/sec )
140
no
60
60
30
30
30
40
50
60
50
40
70
30
50
40
30
60
70
80
W/F
(g-sec/cc)
0.15
0.15
0.10
0.00
0.283
0.283
0.283
0.212
0.170
0.142
0.170
0.212
0,121
0.283
0.170
0.212
0.283
0.142
0.121
0.106
W/F
(g-sec/
mole S02)
1.07
1.07
0.787
0.81
2.03
2.16
2.00
1.85
1.48
1.17
1.38
1.77
0.93
2.45
I1. 50
1.68
2.29
1.19
1.04
0.80
Temp.
(°F)
800
850
850
900
850
850
850
850
850
800
800
800
800
800
900
900
900
900
900
900
%
Conversion
52
68
68
68
83
88
89
98
90
49
59
75
45
90
87
88.
91
85
80
75
119
• MONSANTO RESEARCH CORPORATION •
-------
10
CATALYST B
" T~ T
_jj_
1.0 2.0
W/F, gm-**c/mol« SOj
Figure 48. Effect of W/F on Conversion with
Catalyst-B at 900°F
3.0
120
• MONSANTO RESEARCH CORPORATION •
-------
100i
CATALYST B
1.0
1
2.0
3.0
W/F, gm-iec/mole SOj xlO~
Figure 49.. Effect of W/F on Conversion with
Catalyst-B at 850°F
121
• MONSANTO RESEARCH CORPORATION •
-------
10
CATALYST B
1.0
I
2.0
i
3.0
W/F, gm-»«c/mol« SO, x10~
Figure 50.
Effect of W/F on Conversion with
Catalyst-B at 800°F
122
• MONSANTO RESEARCH CORPORATION •
-------
Table 35
CATALYST-B TEST RESULTS
Run
No.
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
% SOp
In feed
(mole)
.300
.292
.302
.286
.325
.302
.293
.285
.301
.302
.322
.305
.309
.295
.302
.307
.304
.314
Catalyst
Weight
(gms)
8.5
8.5.
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
Total
Flow
(cc/sec)
30
40
50
60
70
80
30
40
50
60
70
80
30
40
50
60
70
80
W/P
^ cc '
.283
.213
.170
.142
.121
.106
.283
.213
.170
.142
.121
.106
.283
.213
.170
.142
.121
.106
W/P
/ gin-sec ^
^mole SCV
2.30
1.78
1.38
1.21
.91
.86
2.36
1.82
1.37
1.15
.92
.85
2.24
1.76
1.37
1.13
.97
..82
Temp.
800
800
800
800
800
800
850
850
850
850
850
850
900
900
900
900
900
900
Conver-
sion
80.2
75.3
73.4
69-1
66.5
60.6
89-0
83.5
80.9
73.9
71.5
69.0
89.0
87.0
85.2
82.7
80.0
77.1
123
• MONSANTO RESEARCH CORPORATION •
-------
lOO-i
90-
CATALYST E
10
1.0
i
2.0
i
3.0
W/F, gm-s»c/moU SO2
Figure 51. Effect of W/F on Conversion with
Catalyst-E at 900°F
124
• MONSANTO RESEARCH CORPORATION •
-------
CATALYST E
3.0
W/F, gm-sec/mole SO2 X10"
Figure 52. Effect of W/F on Conversion with
Catalyst-E at 850°F
125
• MONSANTO RESEARCH CORPORATION •
-------
lOOi
CATALYST E
3.0
W/F, gm-»ec/moU SOj
Figure 53. Effect of W/F on Conversion with
Catalyst-E at 800°F
126
• MONSANTO RESEARCH CORPORATION •
-------
Table 36
CATALYST-E TEST RESULTS
Run
No.
1
2
3
1
5
6
7
8
9
10
11
12
13
11
15
16
17
18
In feid
(mole)
.31*
.299
.301
.303
.309
.297
.311
.318
2.0
.315
.280
.315
.299
.296
.312
.296
.309
.311
Catalyst
Weight
(gms)
8.5
8.5
8.5
8.5
8.5
8.5
5.0
8.5
. 8.5
8.5
.5.0
5.0
8.5
8.5
8.5
8.5
8.5
8.5
Total
Flow
(cc/sec)
30
50
10
60
30
10
10
50.
60
60
60
80 .
30
10
50
60
70
80
W/F
cc
.283
.170
.213
.112
.283
.213
.125
.170
.112 .
.112
.083
.063
.283
.213
.170
.112
.121
.106
W/F
, gm-sec »
vmole S00
2.20
1.39
1.73
l.ll
2.23
1.75
.97
1.19
.17
1.10
.726
.18
2.31
1.76
1.33
1.17
.96
.82
Temp.
800
800
800
800
850
850
850
850
850
850
850
850
900
900
900
900
900
900
Conver-
sion
65.3
19.2
59.5
10.9
92.7
85.7
72.6
78.6
16.0
77.1
56.0
11.0
91.0
88.7
86.3
81.9
77.0
72.9
127
• MONSANTO RESEARCH CORPORATION •
-------
Table 37
Catalyst
code
Catalyst Typs of Support
Support • trtl
Promoter HattrlH Catalyit
Ota
Plow
tttf Rat* S02
Pronoter oo/tto Cone. <
Convorcion
Efficiency
I
75*F
Mli.l-l
NKb-J
MEC.-l
MEOb-l
MKBC-1
rfECV-2
MtAV-1
MECV-2
KtHV-1
ME'-;V-2
MCrV-2
MENV-1
MlJPV-2
MKKV-6
HENV-?
MENV-;
t5»DIi-10
«E3-2
«EPV-»
MEPV-4
MENV-1
MEPV-2
HECV-2
MESV-1.
ME7V-S
KBLV-1
MENV-1
NSKN-1
MEKV-]
MELV-1
MEKV-1
MEKM-2
MSV-1
KEPV.t
WSPF-1
rtKPV-1
MEBV-1
MEKV-1
SnOj
HaO
Cl-,0,
P»0
SnOj
Cr:0,
BaO
VjO,
VjOj
VjO.
Fuller"o Earth
Fuller's Earth
VjOj.pcjO, Fuller's Earth
V;05 "
VjO,
V20S
v,o,-re2o,
VjOj
V;0j
v,o,
Puller'* Earth
S«02
V,0b
Vj05
v,05
Vj05-Pe,0,
V,0,-PejO,
V,05-Pe,0s
AljOj (Acid)
AljOj (AJll)
Fuller1! Earth
V,05
MnOj
V,05
V,05
V,0S
MnOj
V,0,
v,o,
FejO,
VjOj
Puller's Earth
FiilJer'n Eurth
Fuller's Earth
Fuller's Earth
Fuller's barth
Fuller'« Earth
Puller's Earth
Puller's Earth
Alumina
Puller's Earth
(acid)
Puller'o Lurth
Non«
Sec,
CuSO,
SO.
KSCN
NaSCN
SeOa
KjCO,
K,COj
HaSCN
NaSCN
CuSO,
SeO,
NaSC.'l
L1HSO,
NaSCH
SeO,
KHSO
KHSO
L1HSC
KHSO,
KHSO,
SeO,
K,CO,
P«l(SO,),
KHSO,
BljO,
KHSO,
92
92
92
92
92
92
72
150'F
72
72
72
72
72
200'F
72
72
225'F
72
250'F
72
72
300'F
72
72
72
72
72
o?
9*
350*F
72
72
72
72
72
72
100'F
72
72
72
72
f»50'F
72
72
175'F
72
500'F
72
72
72
72
550*F
72
72
88.
72
8
0
8
4
4
H
4
a
\
. a
a
•a
a
a
8
e
a
8
8
a
8
8
8
a
a
a
8
a
8
a
8
8
8
a
8
8
a
a
a
8
a
D
a
8
a
20
20
20
20
20
20
20
20
?0
20
20
20
20
20
20
to
10
10
40
to
to
40
to
to
to
40
to
40
to
to
1*0
to
to
to
to
40
to
to
to
0.317
O.)12
0.292
0.295
0.293
0.281
0.300
0.301
0.301
0.300
0.303
0.312
0.306
0.320
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
to
to
to
to
to
to
to
40
to
tu
tc
to
to .
to
40
to
to-
to
to
to
0.311
0.301
0.310
0.312
0.324
0.307
0.311
0.319
0.298
0.298
0.308
0.307
0.298
0.298
0.3C3
0.315
0.302
0.297
0.301
0.309
0.316
0.313
0.282
0.798
0.298
.
0.278
0.?8*
C.310
I.t9
J.t<
1.57
1.18
l.tc
0.23
l.Cl
1.60
1.60
1.61
0.22
1.56
1.68
1.6t
1.57
1.69
1.57
1.55
1.62
1.67
0.22
1.35
.
1.66
1.59
1.62
l.ta
1.61
1.77
1.70
0.23
1.66
1.88
1.77
1.85
1.95
I.t8
I.t9
1.91
1.63
.1.63
19.2*
19.?•
lf.l«
14.6*
lt.0«
2.1
22.3
lt.0
B.6
2.0
0.3
1.3
O.b
1.3
1.6
0.9
2.6
2.5
2.2
5:1
0.0
t-i.6
18.1
7.3
5.5
1.0
0.0
7.3
1.9
1.7
1.1
1.3
1.0
1.3
1.3
1.1
.0
o.o
i.t
i.t
1.3
130
• MONSANTO RESEARCH CORPORATION •
-------
Table 37 (Cont'd)
Catalyst
CoJa
Tyca of Support
Fuller1g Earth
Oae
»'» Plow
Support wtl Htf Rate SO,
Promoter Material Catalyst Promotar oe/aac Concf
40 0.323
Convorelon
3.1
600'F
rltC>'-2X3 Cr-Ba-Sn .— - .
MiiPV-'j
.VitHV-S
,*ii-PF-l
Mafv-3
H ^-69- *' ' 1
HliKV-2
MEPV-3
MEBC-2
Mt.RV-lq
MEBC-3
HEKV-12
MERV-l-K
600
514CP-4
"•IERV-1P
MhRV- 1 5'
?t
V205
CuO
Bao
V205
CuO
BaO
v2os
V20,
Pt-Fe
•1,0.,
V205
Pt-Fe
Pt
V20,
CuO
V20,
Pt-Fe
V205
v2os5
V205
V205
V205
ZnO
MoOj
V205
V205
V;0.
V205
V205
V205
V;c5
Fe203
V205
CuO
Sn02
Fe203
V205
v,o5
V205
MoO 3
MoOj
Fe20,
Cr20j
Fc203
CuO
V20,
V:05
V205
WOj
Pe203
ZnO
V205
ZnO
Fe»0 1
MoO,
V20,
Pe203
CuO
CuO
CrjOj
CuO
r>nO 2
v2os
Ta,0,
CuO
Cr203
Vj05
Pe,03
Kn02
V205
V20,
B2ol
Fe203
V26V
V205
Puller's
it
•i
n
Alumina
Puller's
"
n
"
n
n
"
Alunlna
Puller's
fl
Alumina
Alumina
Puller's
n
Alumina
Puller's
' "
"
11
"
"
11
"
"
"
"
n
"
"
"
"
n
n
n
it
"
-7
N
n
n
n
n
™
n
"
n
"
"
11
n
ii
n
n
•
n
"
"
"
n
SK-JJ10
Puller's
Earth .
"
"
n
Earth
n
n
•n
n
n
**
Earth
11
Earth
**
Earth
n
n
it
n
n
n
"
•t
"
"
**
n
it
n
n
n
"
n
n
H
n
0
n
"
ti
n
it
n
n
n
n
n
n
n
n
n
"
n
"
n
"
"
"
Earth
K2C03
KHSOW
KHSO,.
Potash
K2C03
KHSO,.
...V
KHSO,
KHSO,
None
Rb2SO,
KHSO,
C8-Rb-K
Rb2SO,
K2SO,
.. —
RbjSO,
Rb 2SO,
Pe2(SO,,)3
Rb-Cs-K
KHSO,
.._.
Cs-Rb
Potash
KHSO,
Ca2SOv
06,30,
Cs-Rb-K
Potash
Potash
KHSO,
Cs 2SO,
Cs-RbrK
Rb2SO,
C3-R6-K
C8-Rb-K
Pe2(SO,)3
Cs2SOi.
None
/Te2(SO|,)3
K2SO,
....
KHSO,
Cs2SO,
Na2$o,
Ll2SOi.
KHSO,
....
KHSO, '
LlHSOk
Na2beO,
KHSO,
KHSO,
Bl,0)
KHSO,
....
KHSO,
KHSO, '
K2C03
. 72
72
72
72
....
72
72
84
72
64
72
76
.. —
72
72
.
72
64
....
72
72
76
72
72
72
72
72
72
72
72
72
72
72
72
72
84
80
72
72
84
80
64
84
72
72
84
84
84
84
72
64
84
64
7;
72
64
84
72
88
72
84
64
72
72
8
8
8
8
.._.
8
6
8
8
8
8
8
8
8
8
8
__ — .
8
8
4
8
8
8
8
8
8
8
8
8
8
6
4
4
8
8
20
8
8
8
8
20
8
8
8
8
8
8
8
8-
D
8
8
8
8
8
8
a
8
'8
.8
8
8
8
8
8
8
8
8 '
10
10
10
20
10
10
20
20
None
20
20
20
8
8
20
20
....
20
20
• _..
20
20
20
20
20
20
20 .
20
20
20
20
20
20
20
20
20
Hone
-.__-
20
20
....
20
20
20
— -
20
20
?0
20
20
20
....
4
20
20
20
20 '
40
40
40
40
40
40
40
.40
40
40
40
40
40
' 40
10
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
• 40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
0.301
0.316
0.316
0-303
0.309
0.301
0.298
0.309
0.307
;0.315
0.307
• 323
.288
0.303
0.31?
0.313
0.301
0.299
0.313
0.317
0.303
• 334
0.295
0.301
0.334
0.334
0-335
0.301
0.303
0.299
0.318
0.334
0.312
0.319
0.301
0.300
0.331
0.323
0.307
0.305
0.312
0.300
0.308
0.293
0.295
0.324
0.321
0.297
0.302
0.317
0.302
0.324
0.290
0.294
0.319
0.295
0.319
0.293
O.J93
0.296
. 0.292
0.295
0.315
0.292
0.309
1.62
1.87
1.87
1.52
1.31
1.64
1.74
1-31
1.59
1.84
1.71
1.41
1.58-
1.53' .
1 .86
1.85
1.22
. 1.29
1.45
1.50 .
1.05
1.36
1.59
1.56
1.53
l.:J6
1.62
1.53
1.69
1.46
1.36
1.32
1.42
1.51
1.67
1-75.
1.47
1.49
1.41
1.55
1.60
•1.51
1.72
1.54
1.79
1.38
1.58
1.60
1.59
1.63
1.29
1.71
1.69
1.73
1.89
1.46
1.80
1.64
1.69
1.59
1.96
1.34
1.78
1.72
92.0*
75.1
46.2
38.6*
34.3
34.2*
30.2
20.1*
20.0*
15.1*
15. 0»
10.8
9.0
8.9
8.6
8.5.
8.0
8.0
7.2
6.9
6.9
6.5
6.4
6.3
5-7
5-1
5.1
5.1
5.0
5.0
4.6
4.5
. 4.5
4.4
4.3
4.0
•3.9
3.7
3-6
3-6
3-6
3-3
3.2
3-1
3.1.
3-1
3.1
3.0
3-0
2.8
2.6
2.5
2.3
2.0
'1.9
1.7
1.6
1.4
1.4
1.4
1.0
1.0
0.0
0.0
0.0
131
• MONSANTO RESEARCH CORPORATION •
-------
Table 37 (Cont'd)
Cui.alyst
Coilo Cotulyjt
'•liiCI'-J
KLFV- 1
IIP-1
MERV-l-K-600
MENV-1
511-CP-1
MEKV-12
MEKV- 7-600
MECV-3-K-600
MEKV- 12-600
jll-CP-3
MEKV-8-600
511-C?-2
MEKV-11-600
MKKV-9-600
MEKV-10-600
MEFP-1
MEPV-1
MEPV-3
MEST-I
MEKV-1
MEPV-1
MEKV-2
MEKV-1 '
KELV-1
CiEyv - 1
MEKV- 'J
MEMV-1
MEKV- 3
MEP-1
MEZV-1
NKM-1
ME KM-?
MS2P-J
MEMP-1
HEPV-5
MEWV-1
ME9V-1
MEKV-5
H£BV-:
MEV-1
(TKKM-1
KKCV-3
HlilV-U
Ol'lCI-1
H<-
r't t,,*p 1
MUNP-1
514CF-2
514CP-3
NKKV-5
MEPP-1
MLPV-4
ME! V-l
MKPV-3
MfcKV-3
MEKV-1
MiCKV-4
MfcKV-1
MELV- I
MEKV- 5
MKZP-1
MF.P- 1
MLMF- 1
MfcUV- 1
MEKH-2
I'l-Pc
Pt
V,0,
Pt
V20S
Pt
Pt
V,05
V20S
V20,
VjO,
Pt
V205
Pt
v,o,
vloj
Pe,0j
V205
Pe,0,
Ta'jO,
vjoj
v|oj
V,0,
VjOj
Pe,0,
vao.
MoO)
pj,3,
V205
ZnO
MoO,
Mn02
Pe20,
ZnO
Pe,0,
MoOj
V20,
WO 3*
V20,
vjot
B20,
V205
KnOj
VjO.
Pt
Pe
rt
Pt
Pr
rt
PI
pt
V;,05
v'jGs'
vCo°'
VJOj
V;0 J
v,ob
viol
V,05
v,o5!
V,05
FejOJ
ZnO
Ke 20 '
MoO]
WnT
V Oe
•/nO
Type of Support
Alumina
Alumina
Puller's
Alumina
Puller's
Alumina
Puller's
Puller's
Puller's
(I
It
''
«
"
It
"
"
11
n
n
II
"
"
Alumina
Puller's
n
n
n
n
n
"
"
w
Earth
Earth
Earth
n
n
Earth
Earth
n
n
n
"
«
w .
H
It
n
"
n
it
;
n
"
Earth
11
n
n
n
"
"
"
Puller's Earth
Al?0j 'acid )
Alumina
„
n
n
Fuller' a
11
n
n
it
„
ii
"
»
„
"
n
Alumina
Puller* 3
Earth
"
» '
"
„
"
I
«
„
«
n
Earth
Support xtl
Promoter Material Catalyst
».
Potash
Rb-K
NaSCN
Ce-Rb-K
Cg-Rb-K
ca-K
C8-Rb-K
CB-Hb-K
C0-Rb-K
Cs-Rb-K
C,-Rb-K
PejtSOJj
Pej(SOi)j
• Na»SeO»
KH30k
Potash
tiy30 1,
None
None
None
None
None
Nonf
None
None
B120,
Nor.e
None
KHSO^
CCjSO,
K2CO,
P«2(SO»))
NaSCN
KHSO,.
?ej(SO, >j
Potash
KhSOu
KliSO,
KHSO,
KIISO.
KHSO,,
L1HSO,
None
None
None
None
BIjCj
None
650'F
..-
72
700'F
z
72
72
72
72
72
72
72
72
72
7S
72
72
76
72
74
72
72
72
64
72
80
81
80
72
81
81
81
81
88
72
81
92
72
750'F
7?
72
800'F
__._
72
72
. 72
72
72
74
72
72
76
•11
72
81
81
80
84
88
72
84
a
8
8
8
8
8
8~
"e"
8
8
8
4
4
8
8
4
1
6
8
e
4
^
a
8
e
20
8
8
20
8
8
8
8
8
8
a
8
a
a
8
8
8
_-._
„_
8
8
4
a
8
6
8
e
4
4
4
8
8
8
8
8
20
3
8
8
8
8
8
Das
Clou
wtS Kate
Promoter oc/sec
—
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
20
....
20
---_
20
4
20
---.
20
2
2
•
20
20
20
20
20
iO
20
20
20
?P
20
4
20
?0
10.
no :
40
10
10
40
40
10
10
40
10
10
40
10
10
10
10
10
10
40
40
10
40
40
10
40
10
40
40
10
40
10
10
10
10
10
10
HO
10
40
40
40
40
40
10
40
"0
10
40
40
10
40
40
10
40
40
40
4U
40
40
40
40
40
'"0
40
40
40
40
40
SO]
cone. %
0.300
0.289
0.315
0.310
0.311
0.303
0.288
o'nt,
0.311
0.313
0.287
C.317
0.291
0.3H
0.327
0.331
0.312
0.321
0.310
0.288
0.306
0.300
0.320
0.312
0.307
0.311
0.310
0.302
0.307
0.305
0.301
0.299
0.312
0.309
0.311
0.303
0.302
0 . 309
0.316
0.305
0.301
0.297
0.298
0.290
0.2')5
0.292
0.306
0.305
0.296
0.308
0.309
0.324
0.291
0.313
0.202
0.301
0.289
O.J89
0.289
0.293
0.301
0.319
0.292
0.306
0.317
0.313
0.^16
0.302
Conversion
Efficiency
W/P t
2.30
1.77
1.49
1.11
1.33
1.91
1.61
1.11
1-31
1.32
1.45
1.28
1.31
1.08
1.33
1.39
1.36
1.60
1.56
1.39
1.61
1.42
1.56
1.62
1.65
1.62
1.70
1.55
1.55
1.19
1.66
i.53
1.69
1.56
1.60
1.55
1.55
1.51
1.72
1.35
1.42
1.66
1.57
1.49
1.7!)
1.01
1.32
1.91
1.52
1.07
1.19
1.72
1.51
1.72
i.50
1.48
\'.6?
L .00
I. SO
1.81
i .',:
1.66
l.so
l.Sfr
l.lt .
J.df)
1.17
10.7
37.0
3.8
33.5
16.5
15-5
12.8
11.2
10.7
10.6
10.3
8 7
8.5
8.2
6.9
6.6
6.5
6.1
5.6
5.5
5.2
l.l
3.5
•3-1
2.9
2.6
2.3
2.3
2.3
2.3
2.0
2.0
2.0
2.0
1.9
1.9
. 1.6
1.3
1.3
1.0
0.9
0.7
0.3
19.5'
15.1
80.3
6Q.4
63.0
60.1
27.9
27.0
26.3
10.7
E.6
7.9«
7.4
6.5
6.3
6.?
5.9
1.9
4.1
4.0
J-5
3.1
2.6
2.5
1.9
1.9
1.7
132
• MONSANTO RESEARCH CORPORATION •
-------
Table 37' (Cont'd)
Catalyst
Code
Type of Support
vt*
Support
Promoter Material
Conversion
Efficiency
it
MliWV-l
MEV-1
MEKM-1
MEM-1
MEBV-2
MEHV-1
V205
WO,
V205
Mn02
MoOj
V205
B203
V20S
Mod 3
None
None
XHSO,,
None
None
None
81
92
12
80
81
81
20
20
20
10
10
10
10
10
10
0.301
0.320
0.301
0.321
0.295
0.300
1.56
.36
65
,11
15
1.60
1.6
1.6
1.5
1.1
1.3
850°F
MEV-l
V205
Fuller's Earth
92
10
0.315
1.38
10.5
900°?
illiPV-5
HP-1
MEPV-1-500
MENP-1
511CP-3
511CP-2
HI-:PV-S
MEPV-1
MERV-1-600
MtiRV-le
MECV-3a
MERV-lf
5HCP-1
MERV-lg
MliCV-3b
MECV-3
MERV-1-100
HLPV-3
HECV-3-100
MEPV-1
MECV-3-500
KEHV-1-500
HEPZ-1
MLRV-1
MECV-3C
WEPM-1
MEPV-3
MECV-3-600
MECV-3
HF.RV-1
MEKV-5
MEPV-ld
HERV-1-500
MEPZ-J
HEKV-j
Hf.KV-2
MliKV-1
MEPV-1
MUFF-1
KtiFV-1
HEf.V-1
Mlil.'V- 3
MEFV-3
ML.LV-1
HEFV- 1
KELV-2
MhK-1
liiPV-S
MkiZF-1
MEMP-1
MKKM-1
MliWV- 1
MKHV-1
MFV-1
HEUV-1
MEZV-1
NEKM-2
KEM-1
MEUV-;;
V205
Pt
v2o5
pt
rt
Pt
V20S
V205
V205
V20S
V205
V205
pt
v2o5
V20S
V205
V205
V205
V205
V205
V?05
v2os
ZnO
V205
V205
MoOj
v,os
V205
V205
V205
V205
V205
V205
ZnO
V205
Vj05
v2o5
v2o5
Fe203
V205
Fe203
V205
V205
v2os
v2os
V20S
Fe203
V205
Fc203
V20S
Fe203
Fe20j
ZnO
Fe20j
Mo03
MnO,
V20,
WO,
v2o5
Mo03
V205
v2o5
v2o.
ZnO
MnO,
MoOj
V20b
B203
Al203(acld)
Al203(acid)
Alumina
— --
.___
Al203(acld)
H
Fuller's Earth
n ti
n n
Puller's Earth
n 11
n ti .
n ii
n n
n tt
Al203(acld)
Puller's Earth
n ti
n ii
n n
0 II
n ti
ii n
n n
ii «
SK-IUO
Ai2o3(acid)
Fuller's Earth
ii it
n it
» n
n n
n n
. u n
" "
n ii
n n
n rt
n »
tt 11
n . n
n ii
rr ir
n ' u
..
u ii
tt n
tr n
1.
II II
II il
Alumina
Fuller's fc'arth
KHSO.*
pp. / Cf), \ .
r e2 \ oui^ ; 3
KjCOj
NaSCN
.
• KHSO,.
K2COj
Rb2SOi.
CsjSO,.
Rb2SOi.
Cs2SOi,
Cs2SOu
Rb2SO,,
KHSO,,
Cs2SOi,
K2C03
Cs250i,
Rb2SO,.
Potash
Rb2SOi»
Cs2SO,,
Potash
KHSO,,
Cs2SOi,
Cg-SOl.
Rb2SOi,
KHSO,,
K2C03
Rb2SOk
Potash
KHSO,.
KHSOi,
KHSO,,
Potash
F«!(So!j|
KHSOt,
Ha2SO,.
Fe2(SO,,)3
LiHSO,.
KHSOi.
Ll2SOi.
None
None
None
None
KHSO,.
None
None
None
B120,
None
KH.SOi,
None
None
72
72
• _— --
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
71
76
72
72
72
72
72
72
72
72
72
80
81
81
81
72
81
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8
-—_-
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20
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„
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.10
10
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0.320
0 2QS
V * cy?
0.293
0.291
0.302
0.303
0.320
0.279
0.298
0.291 •
0.287
0.296
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0.286
0.295
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0.296
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1.81
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1.53
1.99
1.21
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1.81
1.60
1.95
1.98
2.02
1.96
1.51
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1.58
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1.37
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1.96
1.62
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1.55
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1.62
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1.67
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53.3
53.0"
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2.7
1.7
1.3
133
MONSANTO RESEARCH CORPORATION
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EFFECT OF CATALYST PARTICLE GEOMETRY ON PRESSURE DROP IN A FIXED BED
In selecting the geometrical form for catalyst particles in a fixed-
bed reactor, several factors require balancing. It is desired to
minimize pressure drop through the bed within the limits allowed by
contact time requirement and gas velocity, minimize catalyst attrition
rate, and minimize catalyst cost (i.e., not increase cost due to
unusual geometry).
A computer program was written to calculate pressure drop as an aid
in the design of the catalytic reactor. Three equations, were included
in the program based on the magnitude of the Reynolds number. For
laminar flow through the catalyst bed (Re <10.0) the equation becomes
(Ref. 1).
AP - 200.0 GyLA2 (l.Q-0)2
&* - g Dp< p 6'
The transient flow (Re 10.0- 200.0) equation is (Ref. 96):
2.0 f Gn LX3-n (1.0-S)3"n (2)
D(3-n> g p 63
where f and n are evaluated from a least squares fit of two graphs
(Ref. 2). of Re versus f and Re versus n.
The turbulent flow (Re > 200.0) equation is (Ref. 3.),.
AP = "-vg-o G1'9 u0'1 X1'1 (1.0-6) (3)
g Dp1'1 p 63
Each of the above is then converted to psi/ft by:
Ap = P/11U.O
136
• MONSANTO RESEARCH CORPORATION •
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NOTATION
A = surface area of a pellet (sq ft)
D_ = effective particle diameter (ft)
G = mass velocity of gas (Ib/sq ft hr)
L = length of bed (ft)
Re = modified Reynolds number (Dp up/y)
V = volume of a pellet (cu.ft)
f = modified friction factor
g = acceleration due to gravity (4.18 x 108 ft/hr2)
n = state of flow factor
u = linear velocity of fluid (ft/sec)
6 = fractional voids in bed
AP = pressure drop (Ib/sq in/ft)
X = shape factor
y = fluid viscosity (Ib/hr/ft)
p = fluid density (Ib/cu ft)
For the catalytic reactor of this report, all Reynolds numbers were
in the turbulent flow range and equation (3) was used.
Three shapes were considered, those being, cylindrical, spherical,
and granular. For the cylindrical case, the area and volume were
found for three lengths and diameters were L/D = 1.0 (D = 1/8,
2/8, 3/8 in.) and the following calculations were made:
(Ref. 1)
Dp = (6VA)1/3
137
• MONSANTO RESEARCH CORPORATION •
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6 = 0.33
For the spherical case X = 1.0; Dp = 1/8, 2/8, 3/8 In.; and 6 <= 0.36.
For the granular case X = 1.0/0.7 (Ref. §8); 6 = 0.4; and DD = O.i5,
0.35, 0.525 in.
Figure 5^ summarizes the results of the .study of the effect of catalyst
particle geometry on pressure drop. Based on these comparative results,
the suggested catalyst geometry is a cylinder about 3/8 inch in diameter
with a length-to-diameter (L/D) ratio of. 1 or 2. This is a common
form for commercial catalysts.. A spherical shape would result in lower
pressure drop for equivalent, diameter, but it is more costly to. pro-
duce because of its shape. The granular shape, with still lower Ap
response, is subject to higher attrition rate.
138
• MONSANTO RESEARCH CORPORATION •
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200
160-
2
CD
U.
O
I
u
c
I
O
£
Ul
126
Cylinder I/I
C/llmUr 2/1
Cylinder I/I
QruwUr e.li
OmwUr o.J
100 -
25
10 16 20
FLUE GAS VELOCITY, U, Ft/toe
Figure 54. Effect of Catalyst Particle Geometry
on Pressure Drop in a Fixed Bed
139
• MONSANTO RESEARCH CORPORATION •
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Table 38
»10CS(CARDtll3? PRINTER)
• ONF WORD 1NTFGF.RS
•LIST SOURCE PROGRAM
C
~C" ~"PROfiR"RM"TO~OnLCUT,ATE~ MKfSSURt DKUP IN A CATALYTIC BED.
C PROGRAM IS IN FORTRAN-IV FOR THE IBM 1130 BK WITH DISK.
C THE EQUATION USED DEPENDS ON WHETHER THE GAS FLOW IS
C LAMINEHt TRANSIENTt OR TURBULENT.
C THE REQUIRED PARAMETERS FOR THE EQUATIONS ARE-
C EFFECTIVE DlAv-ETER OF THE PELLET(DIAP)
"C THE SHAPE"Mt IUK ur Iht HtLLti(ELMDA)
THE VOID SPACE OF THE REACTOR USING THESE PELLETS (DELTA)
THE GAS DENSITY (RHOI
THE r,AS VISCOSITY (EMU I
THE VELOCITY OF THE GAS '(U)
DP IS PRESSURE DROP IN LB./ SO. IN./ FT.AND IS PER UNIT
WFBSuxE "i>ii»T "f"tk'""K
DIMENSION TITLEI3.6I
SG=ft.l
AL'UO
DO 1 1=1.3 " '
1 READ12 .1010) (T 1TLE I I iJ) «J=1.6)
1010 FORMAT (6A* I .......
C
------- -on -TO- j^rrj ---- ........ ------------------- : -
WRITFIl.tlOXTITLEUtKI.K-l .6)
10 FORMATllOX. 'PARAMETRIC STUDY OF PRESSURE DROP THROUGH CATALYTIC BE
10
(70 TO |20»21» 22>tJ
C DEFINE THE REOUIRFD PARAMETERS FOR THE CYLINDRICAL CASE
C • ' '
20 CONTINUF
OELTA=0.33 • ......................
HNa( 1.0/a. 01/12. 0
TIAP=(6.0*V/3.142)«»0.3331
El.MDA»0.205«A/V»»0.6667
GO TO 23
DEFINE THE REQUIRED PARAMETERS FOR THE SPHERICAL CASE
21 CONTINUE
ELMDA=1.0
DELTA=0.36
GO TO 23
C .
C DEFINE THE REQUIRED PARAMETERS FOR THE GRANULAR CASE
C • ......
22 CONTINUE
-- OCLTA-0** -
OIAP=0.15/12.0
140
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Table 38 (Cont'd)
c
23 WRITEI3.1111>
1111 FORMAT(lOX,'EFF OTA I FT I1.' LTWDJT^ ',* DtlTA
WRITEI3.1110)I>IAP,ELMDA.DELTA
1110 FORMAT! 10X»E12.5»1X,F7.4,6X»F7.4)
C
C DEFINE GAS DENSITY AND VISCOSITY
C
RHO«U.U3Z5
F.MU=0.0798
WRITE(3«1211)RHO
1211 FORMATdOXt'DENSHY OF THE GAS IS1.Ell.4,' LBS/CU FT')
WR1TE(3,1212IEI»!U
FORMATUOX.'VISCOSITY OF THE GAS IS',Eli.4,« LBS/HR/FT'/)
DEFINE GAS VELOCITY
CALCULATE THE MASS FLOWTGl
C i
G=RHO»U»3600.0
WRITE(3.1001)U
1001 FORMAT1 lOX.'GAS VELOCITY (U) ISSFS.Of' FT/S'E'C1)'
C .
~C CTCfULAIt RtYNOCDS IMUMBEH [RE i
FVU »3600.
C
C IF REYNOLDS NUMBER IS LESS THAN 10.0 USE THE FIRST EQUATION.'
C _ IF RFYNOLOS NUMBER IS GREATER THAN 10.0 BUT LESS THEN 200.J
"C~" CISE" THE bErDHn~tOUATION.
C IF REYNOLDS NUMBER IS GREATER THAN 200*0 USE THE THIRD EQUATION
C .......
IFIRE-10.0130*40,40
30 CONTINUE
WRITE I 1 • lUlllRt
100 FORMATf 10X, 'LAMINAR FLOW INDICATED BY REYNOLDS NUMBER OF'tE12.5l
OP»200.0»G»EMU«AL«ELMDA«»2»(1.0-OELTA)»*2 ...........
DP»DP/ISG»OIAP»»2»RHO»DELTA«»3I
OP=DP/144.0 ...........
r;TT TTJ""90
C
40 IF(RE-200.0)50.50i60 "' .........
C
50 CONTINUE ............
WRITE! 3,2 00 1 RE _ _ _ _
ZOO" I- UKlAll rtJH", • I KAIN 5 1 TTDrn«L hLUW INDICAltD IJY KEYNOLD5 NUMBER QF'«E
112. SI
REMOD=ALOGTIRE> " ........
EN-0.522»REMOD*0.4B187
F=10.0»»J-0.62«REMOD+1.6428) . ..........
OP«2.0«F*G««EN«AL«ELMOA«»(3«0-ENI«(1.0-OELTAI**<3.0-EN»
,,^-(J(l/-|uJnt,ww | 3,u- tIM I w3
OP = DP/144.0
GO TO 90
141
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Table 38 (Cont'd)
c
60 CONTINUE
300 FOHMATI 10X . MUKBULENT FLOW INDICATED BY REYNOLDS NUMBER OF'.EIZ.
15)
(>P»0.02«.3»G»»1.9«EMU»»0.1»ELMDA»«1.1«I 1.0-DELTA)
nP=nP/tDIAP»»i.l«SG»RHO«DELTA«»3)
90 CONTINUE
• WRITE(3idOO)DP
40C FORMAT) 10X, 'OEL P«'tE13.5t' PSI/FT'/I
C
U=U+10.0
80 CONTT17OT
WRITEO.11121
1112 FORMAT I MM
70 CONTINUE
CALL EXIT
END
112
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APPENDIX III REFERENCES
1. Leva, M. , et al, "Introduction to Fluidization," Chemical
Engineering Progress, Vol. M, No. 7, p. 518.
2. Ibid, p. 519-
3. Foust, A. S., et al, "Principles of Unit Operations," J. Wiley
and Sons, Inc., New York, 1962, p. ^76.
• MONSANTO RESEARCH CORPORATION •
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APPENDIX IV
REACTOR DESIGN
Preceding page blank
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APPENDIX IV
REACTOR DESIGN
Optimum design of the catalytic reactor was based on minimum total
capita- and operating.costs, which are functions of the costs of the
reactor, fan, and power calculated on a yearly basis.
Total Cost of Reactor = annual cost of (reactor+fan+power)
The total catalyst .volume was fixed by the. reaction rate, determined
experimentally, for SQ% conversion of sulfur dioxide to sulfur tri-
oxide. Consequently, for a given reactor diameter, the height of
the catalyst bed is prescribed. It then becomes a matter of determin-
ing the most economic geometry for the requisite catalyst volume.
If, for example, a given reactor diameter, or cross sectional area
of the catalyst bed is increased, the installed cost of the reactor
increases; but, due to a corresponding reduction in bed height, the
pressure drop decreases, reducing the cost of fan and power. An
economic balance shows that at. one particular equivalent 'diameter,
the reactor cost is minimum. An increase beyond this point raises
cost of the reactor much faster than it lowers the cost of fan and
power. Hence, this represents the point of optimum reactor design.
The assumption employed in the optimization program for the large
new power plant facility were:
1) Flue gas rate: 2.5 MM SCFM
2) Flue gas temperature: 850°F
3) Total volume of catalyst: 20,000 ft3
4) Power cost: $0.006/KWH
Installed cost of reactor: $2,500/ft (for 30-ft diameter)
5)
6)
Installed cost of fan in $: 85.8 (x°<699), where x = drive
horsepower
7) Depreciation rate for reactor and fan: 10 years, straight
line depreciation
The design parameters for .the reactor installation for a 1400 MW sta-
tion, based on our optimization study are:
1) Number of reactors in parallel: 4
2) Number of .catalyst beds in parallel in each reactor: 13
1116
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3) Height of each bed: 6-1/1 in.
4) Height of each segment: 9 ft.
5) Overall height of each reactor: 117 ft.
6) Diameter of each reactor: 30 ft.
7) Size and shape of catalyst: 3/8 in. cylinders
8) Total pressure drop.: 0.96 in. H20
9) Total cross sectional area of catalyst: 38,000 sq. ft.
10) Superficial gas-velocity through each bed: 3.22 ft/sec.
The computer program for optimum reactor design is shown in Table 39
1*17
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Table 39
PROGRAM OOPT - CALCULATES THE OPTIMUM DIAMETER OF A REACTOR BED.
REAL KW
DIMENSION ZI20) .INI 10)
PI-3.1M6
INITIALIZATION OF PARAMETERS FOR BED
10 CALL READP(6tZ»IN>
J=IN<1 I
CICWZI 1)
TOTV=Z(2)
D=Z<3)
DD=Z15)
INITIALIZATION OF GAS PARAMETERS
20 CALL READPIbtZtlNI
TF=Z(1)
EMU=Z<2)
GM=Z(3)
AMWZKtl
P-ZI5)
TR'TF+460.
RHO-AMW«P/(10.T2«TRI
1-0
PRINT INITIAL PARAMETERS
30 WRITE! 3.600) J.CHW.TOTV.D. DO. THETA.RHO>EMUtGM(TF»AMWiP
WR1TEI3.510)
35
H"TOTV/XS£Ct
GUX»GM/XSECT
CALCULATE THE COST OF THE REACTOR
VS=PI*15.»»2.«H
SEGM=A1NT ( TOTV/VSI+1.
1F(H-8.0)38O8.36
36 CSTR=14.3«30.»H»SEGM
GO TO 39
38 CSTR-2000.»SEGM
CALCULATION OF PRESSURE DROP
39 CALL PDROPI J.RHO.EMU.GUX.RE.DP)
CALCULATE COSTS OF POWER AND FAN AND TOTS
DAP=2.602E<.«DP
CSTF=85.8»DAP«»0.699
CSTP=THETA»CKW»KW
CSTOT-CSTR+CSTF+ CSTP
WRITEI3t520)XSECT.SEGM,CSTFtCSTP.CSTR,CSTOTi H tDP
<»0 CALL DATSWU.IO
GO TO 160.50) ,K
50 0«D+OD
GO TO 35
60 GO TO 10
500 FORMAT (!Hlt3X i 'J=' I 2 <2X> • COST /KWH° ' F9.5 >2X ' TOT . VOL. = ' E12.4 »2X
1 •DIAM.='E12.5.2X1DO»lE12.5.2X'THETA=«E12.5//3Xt1RHO«1E12.S.2X1MU=t
2E12.5.2X.'MASS VEL.« ' F.12.5 .2X» TF = 'F8.2 . 2X 'MW« ' F8.2 »2X« P= • F 10.4/ / )
510 FORMAT(T6'XSECT'T20'SEGMIT341CSTF'T<»9ICSTPIT62>CSTR1T761CSTOT'T90
1' H'TIO^'DELT.P'/I
520 FORMAT(9(2X.E12.5) )
END
• MONSANTO RESEARCH CORPORATION •
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APPENDIX V
RELATIONSHIP OF W/F FACTORS
AND SPACE VELOCITY
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APPENDIX V
RELATIONSHIP OF W/F FACTORS AND SPACE VELOCITY
W = catalyst weight, g
dc = catalyst density, g/liter
Lc = liters of catalyst
v = total gas flow, cc/sec
V = total gas flow, liters/sec = v x 10~3
V = S02 flow, liters/sec
N = moles total gas/sec
n = moles S02/sec
P = total pressure, atm..
T = temperature, °K
R = gas constant = 0.082 liter atm/mole-°K
V
fS02 = fraction S02, mole or volume = -
V
S = space velocity, sec"1 = liters S02/liter catalyst - sec =
V
W
Convert W/F, g-sec/cc total gas flow, Q! = -—
to W/F, g-sec/mole S02, Q2 = -jj
(1) V = v x 10- 3
(2) H,-g-
f PV
(3) n = fSQ2 (N) ,=
for out test system: P = 1 atm
T = 25°C (298°K)
150
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(v)
(0.082)(298) 24.5 x
then
-\ W(2*J.5 x 103) _ 24.5 x 103 Q,
~ —w—ff \ ~
v (fso2)
(PT)
and
(8) S
Convert Q to S.
For the general case:
-"' • -r™H
but W = dcLc
V1
and fS02 = ~V
dcLc(RT) dcLQ(RT) _ /RT
*+)
151
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APPENDIX VI
COMPARATIVE COST OF PLATINUM
AND VANADIA CATALYSTS
Preceding page blank
153
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APPENDIX VI
COMPARATIVE COST OF PLATINUM AND VANADIA CATALYSTS
The literature fails to reveal catalysts for this application better
than vanadla or platinum, which are about equivalent in conversion
efficiency. The possibility cannot be ignored that a better cata-
lyst may be announced at any time in the future. Nevertheless, at
the moment, there are only the two cited. A cost analysis of platinum
and vanadia catalysts summarized, in Table JJO , indicates that the
platinum catalyst must be at least ten times as effective as vanadia
catalyst to be comparable in cost with vanadium catalyst. Decrease
in platinum catalyst cost is not expected in the near future because
platinum and labor costs are increasing.
.154
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Table 40
SUMMARY OF ECONOMICS FOR VANADIA AND PLATINUM CATALYSTS(3)
V Catalyst Ft Catalyst
Volume, Cu ft^1) 30,300 3,020
Initial cost, dollars^2) 1,250,000 12,800,000
Regeneration cost, dollars(2) 1,032,000
Capitalized cost, dollars^2) 3,920,000 20,380,000
NOTES: (1) Estimate of the volume of platinum catalyst that
would be equal in cost to the estimated cost of
vanadia catalyst.
(2) Estimates are based on the same W/F ratio for
platinum and vanadia catalysts.
(3) Assumptions used for comparison of platinum and
vanadia catalysts:
General:
a. rate of return: 8% per year
Vanadia Catalyst:
b. amount used for the oxidation of S02 in flue gas
30,300 ft3
c. price: $4l/ft3
d. packed density: 36.8 lb/ft3
e. catalyst life: 5 years
Platinum Catalyst:
f. catalyst contains 0.5$ platinum
g. packed density: 50 lb/ft3
h. catalyst cost: cost of platinum + $2.75/lb
catalyst, (for manufacturing cost)
i. regeneration cost:. $0.75/lb catalyst
J. regeneration interval: 2 years
k. loss of platinum during regeneration: 2%
1. price of platinum: $110/oz (troy)
m. catalyst life: 30 years (including obsolescence
and/or abandonment)
155
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(1) To estimate the volume of Pt catalyst that would equal
in cost the cost of vanadla catalyst required for a particular
flue gas application.
Initial cost of vanadia catalyst
= 860,000 liters x ^'^
' liter
= $1,250,000
Capitalized cost for vanadia catalyst
= $1,250,000 x
= $1,250,000 x 3-131
= $3,920,000.
Assuming that the amount of Pt catalyst required = P Ib
Initial cost of Pt catalyst, .
A = Cost of Pt -t manufacturing
cost of Pt catalyst
= 0.005 P Ib x ii7|0 , $2^75 x p lb
= 8.8P + 2.75P
= 11.55P
156
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Regeneration cost of Pt catalyst:
H = Refining cost + cost of Pt loss
ID
x P Ib + 0.0001 x P Ib
$1760
= 0.75P + 0.176P
= 0.926P
To calculate the volume of Pt catalyst, capitalized cost of
vanadiu catalyst = capitalized cost of initial.Pt catalyst +
capitalized cost of regenerating Pt catalyst every two years
3,920,000 = AF30 + H (G2-G30)
= 11.55P
0.926P
(1+0.08)30
(1+0.Ob)00
(1+0.08)^-1
(1+0.
= 11.55P x 1.110 + 0.926P (6.010-0.1103)
= 11.55P x 1.110 + 0.926P x 5.8997
= 18.283P
151,000 Ib of Pt catalyst
= 3,020 cu ft of Pt catalyst
157
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(2) at same W/F values of Pt and V catalysts, what weight of
Pt catalyst is required? Cost of Pt catalyst? Cost of V cata-
lyst?
Weight of V catalyst = 30,300 ft3 x 36^.,lb
= 1,115,000 Ib
= 505 x 106 gm
Weight of Pt catalyst = 1,115,000 Ib
S02 = 150 x 106 SCFH x 0.003
= 0.45 x 106 SCFH
10* — x Hr v lb-mole x
xu x sec x x
106 _ ,. 2 mole
H ' L'^ x 1U
136 x 10H ' ' sec
W/F = 505 x 10^ g x sec
mple
-37 x 106 g~sec
'*< x 1U mole S02
Initial cost of V catalyst = $1,250,000
Capitalized cost = $3,920,000
Initial cost of Pt catalyst,
= Cost of Pt + Manufacturing cost of
of Pt catalyst
158
• MONSANTO RESEARCH CORPORATION •
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= 0.005 x 1,115,000 Ib x
x 1,115,000 Ib
= $12,880,000
Regeneration cost of Pt catalyst
= Refining cost + cost of Pt loss
= ^yp- x 1,115,000 Ib + 0.0001 x
1,115,000 Ib x $1?k°
= $1,032,000
Capitalized cost of Pt catalyst
= $12,880,000 x 1.110 x $1,032,000 x 5.8997
= $20,380,000
159
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APPENDIX VII
SORPTION ISOTHERMS FOR MOLECULAR SIEVES
Molecular sieves have recently been shown to be useful in many ap-
plications as heterogeneous catalysts and as vehicles in studies
of catalysts and catalyst mechanisms. However, no publication,
to our knowledge, has revealed specific studies to determine their
merit as support material for known S02 oxidation catalysts.
We obtained samples of several experimental forms of zeolite materi-
als and embarked upon a program to screen the materials as potential
supports. The monodisperse nature of the zeolite pores and their
extremely high specific area are characteristics highly favorable
for maximum contact conditions between the S02 and active catalyst.
Zeolite materials are well known for their adsorption characteristics
with regard to the various constituents in stack gas, and we wished
to know if these characteristics would be an aid or hindrance in our
contemplated use of them as a support material. Initial experiments
were performed at a low temperature range (75° to 250°F) as shown
in Figures 55 through 6l.
162
• MONSANTO RESEARCH CORPORATION •
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901
c
o
a
s Y0
•a
5 50 J
rt
a
a
< 40
30-
SK-20
Bed Wt. =6.9 gms
12 20 28 36 44 52 60 68 76 84
Exposure Time at 40cc/sec Stack Gas Flow, minutes
Figure 55. Sorption for Molecular Sieve SK-20
163
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c
o
o
to
T3
(M
O
CO
c
(b
JH
m
a
a
100
lOO
90-
80-
70-
60-
50
40-
30-
SK-400
Bed Wt. = 6. 7 gms
4 12 20 28 36 44 52 60 68 76 84
Exposure Time at 40cc/sec Stack Gas Flow, minutes
Figure 56. Sorption Isotherms for Molecular Sieve SK-400
• MONSANTO RESEARCH CORPORATION •
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c
0
o
(ft
TD
(M
O
tfl
a
a
100
90
80-
70
60
50
40
30-
SK-^00
Bed Wt.
= 9.5 gms
4 12 20 28 36 44 52 60 68 76
Exposure Time at 40cc/sec Stack Gas Flow, minutes
84
Figure 57. Sorption Isotherms for Molecular Sieve SK-400
165
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c
0
4_J
a
L.
0
90
HO
70
O 60
to
O.
a.
30-
SK-^10
Bed Wt. = 6.2 gms
4 12 20 28 36 44 52 60 ,68 76
Exposure Time at 40cc/sec Stack Gas Flow, minutes
84
Figure 58. Sorption Isotherm for Molecular Sieve SK-410
166
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c
o
o
-------
c
o
0
V)
O
c
rt
a
a
90
80
70
60
50
40
30-
13X
Bed Wt. =8.9 gms
4 1Z 20 28 36 44 52 60 68 76 84
Exposure Time at 40cc/sec Stack Gas Flow, minutes
Figure 60. Sorption Isotherms for Molecular Sieve 13X
168
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c
o
o
(/>
•u
PvJ
O
90
80'
70'
60'
? 50
Q.
a
40-
30.
SK-110
Bed Wt. = 6.4 gms
12 ZO 2« 36 44 52 60 68 76
Exposure Time at 40cc/sec Stack Gas Flow, minutes
84
Figure 61. Sorption Isotherms for Molecular Sieve SK-110
169
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APPENDIX VIII
TYCO MODIFIED CHAMBER PROCESS REPORT
Preceding page blank 171
o MONSANTO RESEARCH CORPORATION •
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APPENDIX VIII
TYCO MODIFIED CHAMBER PROCESS FOR REMOVAL OF SO?
FROM POWER PLANT FLUE GAS *
The basic reactions which describe both the Lead Chamber Process
for sulfuric acid manufacture and the basic Tyco Modified Chamber
Process are shown on page 175« Sulfur dioxide is oxidized with
excess nitrogen dioxide and water to form sulfuric acid and equi-
molar quantities of nitric oxide and nitrogen dioxide (1 and 2)
which are absorbed in sulfuric acid forming nitrosylsulfuric acid
(3). The oxides of nitrogen are recovered by heating the HNS05
after which the NO is reoxidized to N02 CO. In the Tyco Process,
excess N02 is absorbed in water forming nitric acid (5).
The flow sheet in Figure 62 shows the basic difference between the
standard.Lead Chamber Process and the Tyco Modified Process. In the
standard process the NO is present in high enough concentrations to
allow the reoxidation to take place in the same reaction mixture as
the S02 oxidation. In the new process, the NO is too dilute to
allow the slow oxidation to occur in reasonable contact times, thus
requiring the concentration of the NO and separate oxidation.
The process as shown in Figure 62 is profitable, but suffers from
the disadvantage of requiring 31% additional coal to provide the
necessary energy. The heat is needed to (1) vaporize the NO and
N02 for recycling, and (2) to concentrate the acid back to 80$
for recycle.
In trying to overcome these deficiencies, two major breakthroughs
in process concept have been achieved which may well make the
Modified Chamber Process a practical one for removing S02 from
power plant flue gas and may have a profound effect on the sulfuric
acid manufacturing Industry in general.
Figure 63 shows an Isothermal Scrubber in which the gas is scrubbed
with hot sulfuric acid. The entering hot acid has an equilibrium
vapor pressure of water such that all the water that comes in with
the raw flue gas leaves with the clean stack gas. The bottom of
the scrubber is also run at conditions such that there is no net
transfer of water between gas and liquid.
To accelerate the recovery and reoxidation of the oxides of nitrogen
we examined the stepwise reactions on page 178. The third equation
shows that the net reaction is an oxidation of HNS05 which might
well be carried out in the liquid phase. This reaction does not
occur by simply contacting air with HNS05 solution (in P^SO^), but
we have found that by using an activated carbon catalyst, the
reaction does take place. Experimental results comparing charcoal-
"Report presented by Tyco at June 11-13* 1969 Contractors meeting.
1,72
• MONSANTO RESEARCH CORPORATION •
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catalyzed and uncatalyzed reactions are ;shown on page 179. Figure 6^
shows the proposed catalytic stripper which takes advantage of the
greatly accelerated N02 recovery.
These two modifications lead to the simplified Catalytic Chamber
Process shown in Figure 65. The important thing to note is that
no additional heat energy is required above that which comes in
with the flue gas at 300°F.
The sulfuric acid concentration in the recycle (product) stream of
the newly modified system is no longer limited by equilibrium of
vapor pressure considerations as it was in the original Chamber
Process. It is only limited by the loss of HaSOu vapor at the
temperatures required to permit the water vapor to escape with the
stack gas. This places a limit of 92% H2SOi+ on the power plant
flue gas cleaning process, but does not impose the same limit on
the Lead Chamber Process which does not have the problem of water
removal. Acid at 100$ concentration could conceivably be produced
in this simpler Lead Chamber Process using catalytic stripping
techniques.
This removal of the additional heat requirements greatly improves
the economics of the system as shown on pages, 182. andt 183. Use of the
NAPCA guidelines for the catalytic process have raised the total
capital cost and general operating cost estimates to somewhat
higher levels than the more liberal estimates for the baseline
process, but the dramatic difference lies in the saving on heat and
cooling water. This saving amounts to more than $6 million in
annual operating expenses. Using credits shown on page.183, we
estimate the process to earn $2.62 per ton of coal or 39-5? return
on capital investment making the process both technically and
economically attractive.
17.3
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PROCESS REACTIONS
(1) S02 + 2N02 - *• S03 + NO + N02
(2) S03 + H20 - 5-
(3) NO + N02 + 2H2SOi, - »- 2HNS05 + H20
(i|) 2NO + 02 - =- 2N02
(5) 3N02 + H20 - ^ 2HN03 + NO
174
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BASELINE PROCESS
Ho SO,. 80% 80°F
flue gas
• 300°F
2 0.3% SO,
O *
z
en
Z
H
O
rn H- 1
m TO !
JMC
0 VJl tl.
cr\
O po
o
TJ
O
H
O
Z
REACTOR
to stacK -^
330°F
\
H2° COOLER
I
]•» L*MM —
HN03
ABSORBER
product
HNO
3
N203
• I
NO
OXIDIZER
STRIPPER
stripper
2100°F
SCRUBBER
H2S04 76$
HNS05
275°F
I
gas to
stack
r
H,SOU
7 O a
clean
gas —
2100°F
1
I
I
f ACID
> POOT FR
* \yWWJ_f£jJl
^ product
H2S04
ACID
CONCENTRATOR
-------
ISOTHERMAL SCRUBBER
stack gas
250°F
1% H20
60 ppm
N203
flue gas
250°F
1% H20
3000 ppm N203
80*
250°F
.001 M HNS05
H2S0.4 80$
250°F
.11 M HNS05
Figure 63
176
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NITROSYLSULFURIC ACID OXIDATION
i»HNS05 + 2H20 - »- 2N203 + 2H2SOi»
2N203 + 02 - *•
02 + 2H20
17.7
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NITROSYLSULFURIC ACID OXIDATION EXPERIMENTS
Carrier Gas
Nitrogen
Air
Air
Packing
Saddles
Saddles
Charcoal
Concentration (mm Hg)
NO NO,
0.01 0.01
0.01 0.02
0.01 0.75
178
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CATALYTIC STRIPPER
250°F
.11 M HNS05
flue gas
250°F
N02
charcoal
packing
flue gas
250°P
H2SOt, 80$
250°F
.001 M HNS05
Figure
179
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CATALYTIC CHAMBER PROCESS
stack gas
250°F
1% H20
COOLER
REACTOR
flue
gas
1% H20
COOLER N203
HN03
ABSORBER
\ \
gas
100°F
product
HN03 52%
flue gas
300 °P
.3% S02
COOLER
gas
N02
250°F
flue gas
250°F
ISOTHERMAL
SCRUBBER
H2S04
.11 M HNS05
250°F
CATALYTIC
STRIPPER
"*" product
H2SO,,
Figure 65
180
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ECONOMIC ANALYSIS
ITEM
2
o
2
2
H
O
m
(A
m
>
2)
O
I
O
o
3
TJ
O
oo
Major equipment
Total Capital Cost
Annual Operating Cost
Heat
Water
Total
Saving (heat and water)
Other operating costs
Total operating cost
BASELINE PROCESS
$4,535,000
9,2717000
5,220,000
1,133,000
6,353,000
2,159,^00
$8,513,^00
CATALYTIC PROCESS
$ 3,985,000
12,600,000
$6,185,000
168,000
168,000
3,688,900
$ 3,856,900
O
z
-------
BASELINE PROCESS
CATALYTIC PROCESS
o
z
z
H
O
3)
m
(A
m
O
I
O
o
3)
TJ
O
O
Z
oo
Credits
H220k 6 $20/ton (100%)
HN03 % $84/ton
Total credits
Operating costs
Profit
$/ton coal
Return on Investment
$6,201,000
3,530,000
9,731,000
8,513,^00
1,217,600
0.55
13-1?
$6,201,000
3,530,000
9,731,000
3,856,900
5,874,100
2.62
39.5?
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APPENDIX IX
VAPOR PRESSURE APPARATUS AND PROCEDURES
183
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Glass Tubing
2
en
Z
H
O
3)
m
w
m
a
o
i
o
o
X
TJ
O
O
Z
Heater & Controller,
oo
-Cr
100ml Flask
Sulfuric Acid with
Dissolved HNSO,
Rubber Tubing & Pinch Clamps
To Vacuum &Liquid N2&Dry Ice Traps'
u
Magnetic Stirrer
Figure 66
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APPENDIX IX
PROCEDURE FOR OBTAINING VAPOR PRESSURE DATA
1. A quantity of HzSOi* of the desired concentration was weighed
and transferred into the 100 ml flask (approx. 50-60 ml).
2. A quantity of anhydrous nitrosylsulfuric acid was weighed and
transferred to the flask.
Formula: NzOa * (803)2
3. The flask containing the above mixture was cooled in an ice
bath to approx. 2°C.
4. The flask was connected to the rubber tubing and vacuum was
pulled on the system (down to approx. 0.05 mm Hg). This removed
dissolved gases in the acid as well as the gases from the vapor
space.
5. The pinch clamp above the flask was closed tightly.
6. The flask was then placed in the oil bath and heated to the
desired temperature (while stirring).
7. The flask was held at the desired temperature for 1-3 hours.
(Until no evidence of gas evolution was observed).
Q. The remainder of the system was evacuated to approx. 0.01 mm Hg
and the pinch clamp to the vacuum source was closed.
9. The pinch clamp above the flask was opened for about 3 seconds.
This is long enough to equalize the pressure throughout the
system but not long enough for a significant quantity of vapor
to evolve from the solution.
NOTE: The stirring bar was stopped prior to opening the pinch
clamp .
10. The pressure in the manometer was quickly measured and corrected
for the ratio of volumes and the temperature difference. The
two volumes were (1) the free space above the flask up to the
pinch clamp, and (2) the volume from that pinch clamp to the
rest of the system including the manometer leg. The temperature
of the system was taken as the "weighted" average between the
temperature of the oil bath and room temperature.
11. The contribution of water vapor was subtracted from the total
corrected pressure.
185
• MONSANTO RESEARCH CORPORATION •
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12. It was assumed that the total pressure (minus H20 pressure)
was a result of equal moles of NONOa. No association of these
gases or any other gases was assumed.
The step which proved to be the cause of the erroneous results was
step 9- Apparently sufficient vaporization occurs (even with a
quiesent liquid) in two-three seconds to cause a "high" vapor pressure
reading. This was demonstrated by opening the pinch clamp above the
flask and waiting until the entire system was at equilibrium. The
final pressure measured at equilibrium differed by about 25% from the
pressure measured after two-three seconds exposure to the flask.
186
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APPENDIX X
SQ3-PRODUCT ACID RELATIONSHIPS
187
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00
00
HI
a
500
450
400
350
300
250
200
_l
H
[
10
100
1000
SO, IN FLUE GAS, ppm
t
10.000
Figure 6?. Effect of S03 Concentration in Flue Gas
on Dew Point of the Acid
-------
o
2
•z.
H
O
m
o
i
o,
o
o
•a
>
H
O
3D
A,
co
O 'Z)
O 0)
3 M
O P3
(D ct
D H-
ct
H-
O
O I-1
O C
D CD
O.
CD O
D P
• w cn
;pj
jet O
CD CD
S
•n
o
ct
O
O
en
HI
O
Z
o
o
100
90
80
70
60
Cj 50
v>
jjj 40
cc
30
20
10
100 150 200 250 300 350 400
DEW POINT, °F
450
500
550
600
65O
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