APPLICABILITY OF CATALYTIC OXIDATION TO THE
DEVELOPMENT OF NEW PROCESSES FOR REMOVING
SO2 FROM FLUE GASES VOLUME I - LITERATURE
REVIEW
R. E. Opferkuch, et al
NATIONAL 'ECHNICAL INFORMATION SERVICE
Distributed ... Mo foster, serve
and promote the nation's
economic development
and technological
advancement.'
U.S. DEPARTMENT OF COMMERCE
This document has been approved for public release and sale.
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APPLICABILITY OF CATALYTIC OXIDATION
TO THE DEVELOPMENT OF NEW PROCESSES
FOR REMOVING S02 FROM FLUE GASES
Volume I - LITERATURE REVIEW
Contract No. PH 22-68-12
Prepared by
R.E. Opferkuch
S.M. Mehta
A.H. Konstam
D.L. Zanders
H.R. Strop
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
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BIBLIOGRAPHIC DATA
SHEET
1. Report No
Report No.
APTD-0675
3. Recipient's Accession No.
4. Title and Subtitle
Applicability of Catalytic Oxidation to the Develop-
ment of New Processes for Removing S02 From Flue
Gases Volume I - Literature Review
5- Report Date
August 1970
6.
~E~!Opferkuch,Project Leader, S~7FT Mehta,
7. Author(i) *» •
A. H. Konstam, D. L. Zanders, H. R. Strop
8. Performing Organization Kept.
No.
9. Performing Organization Name and Address
Monsanto Research Corporation
Dayton Laboratory
Dayton, Ohio 45407
10. Project/Tnsk/Wotlc Unit No.
11. Contract/Grant No.
PH 22-68-12
12. 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 & Period
Covered
14.
IS. Supplementary Notes
16. Abstracts
An extensive literature search for pertinent information relative to the
catalytic oxidation of sulfur dioxide, is presented lit this report.
This, the first phase of - th«-program., also attempts to identify; describe
and evaluate processes, disclosed in the literature to have commercial
potential for removal of sulfur dioxide from flue gas by oxidation.
17. Key Words and Document Analysis. 17o. Descriptors
Catalysis
Reaction kinetics
Oxidation
Sulfur dioxide
Ca talys ts
Vanadium
Magnesium Oxides
Electric Power plants
Flue gases
17b- Identifiers/Open-Ended Terms
Kiyoura proce&s
Montsanto process
Mitsubishi process
Tyco process
I7e. COSATI Field/Group
13/B
18. Availability Statement
Unlimited
19.. Security Class (This
Report)
UNCLASSIFIED
20. Security Class (This
Page
UNCLASSIFIED
21. No. of Pages
239
22. Price
FORM NTI«-(B (1O-70)
USCOMM-OC 40J29-P7I
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MRC-DA-245
APPLICABILITY OF CATALYTIC OXIDATION TO
THE DEVELOPMENT OF NEW PROCESSES FOR
REMOVING SO2 FROM FLUE GASES
Volume I - Literature Review
Contract No. PH 22-68-12
MRC Job No. 6708
Prepared by
R. E. Opferkuch, Project Leader
S. M. Mehta
A. H. Konstam
D. L. Zanders
H. R. Strop
MONSANTO RESEARCH CORPORATION
DAYTON LABORATORY
Dayton, Ohio
August 1970
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 ^5277
-------
FOREWORD
The intent of this volume is to present in an organized manner
the accumulation and assessment of all available literature
data on S02 removal from flue gases by new and existing cata-
lytic oxidation processes as of 31 December 1968. In that
this volume represents the first phase of a three phase pro-
gram of concurrent tenure, occasional reference to phases
two and three will 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 labora-
tory 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.
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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 infor-
mation relative to the catalytic oxidation of sulfur
dioxide, i.e., active materials, mechanisms of catal-
ysis, 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 ap-
plication to removal of sulfur dioxide from flue gas
by catalytic oxidation.
c. Identify at least one effective catalyst for the de-
sired application and design a process for removal
of sulfur dioxide from flue gas by catalytic oxida-
tion 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 practi-
cal gaseous catalyst noted.
b. Kinetic equations describing conversions of sulfur
dioxide over vanadia or platinum catalysts were de-
rived 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 concen-
trations of sulfur dioxide found in flue gas.
c. A number of processes were described as having com-
mercial potential for flue gas cleaning. Compara-
tive 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.
ill
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d. The most practical mode of recovery of oxidized sulfur
value from a use standpoint, in this country, is produc-
tion of fertilizer grade sulfuric acid.
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 effect-
ive as vanadia, platinum or nitrogen dioxide in convert-
ing sulfur dioxide to trioxide.
d. Nitrogen dioxide was the only practical "low temperature"
catalyst observed.
e. Using nitrogen dioxide as a catalyst, it is potentially
practical to remove both sulfur dioxide and indigenous
nitrogen oxides from flue gas simultaneously.
From preliminary process designs and cost estimates, bas.ed 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 vanadia based processes are likely
to be in the range of 0.50 to 0.7^ mllls/Kw-Hr generated
before sulfur value net back.
c. A substantial reduction in capital and operating costs
are potentially available through a technique of sorb-
ing 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.
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VOLUME 1
TABLE OF CONTENTS
Page
1. SUMMARY 1
II. CONCLUSIONS 4
III. DISCUSSION 6
A. INTRODUCTION 6
B. CATALYSIS IN S02 OXIDATION 8
1. General Remarks 8
2. Factors Affecting Activity of Solid 9
Catalysts
3- Comparison of Platinum and Vanadium 11
as Catalysts
l*. Conversion of S02 Under Equilibrium 14
Conditions
5. Conversion of S02 Under Non-Equilibrium 17
Conditions
a. Introduction 17
b. Oxidation Over Vanadium Catalysts 18
c. ' Oxidation Over Platinum Catalyst 35
d. Oxidation of S02 on Chromium Oxide 38
Catalyst
e. Oxidatioh Over Iron Oxide Catalysts 39
6. Catalyst Poisoning and Effects of Inert 40
Gases
7. Oxidation of S02 in an Electrical 42
Discharge
8. Feasibility of Photochemical and Radia- 43
tion Methods for S02 Oxidation
a. Introduction 43
b. Direct Irradiation of S02-Containing 43
Flue Gases
c. Irradiation Pre-Treatment of Catalysts 47
for S02-0xidation
9. Irradiation of S02 Containing Flue Gases in 49
the Presence of Catalysts
10. Liquid Phase Catalysis 49
11. Homogeneous Gas Phase Catalysis 49
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TABLE OF CONTENTS (Cont'd)
Page
APPLIED CATALYSIS IN FLUE GAS TREATMENT 50
1. General 50
2. Vanadia Based Processes 51
a. Monsanto-Penelec Process 51
b. Kiyoura-T.I.T. Process 5*1
c. Bayer Double Contact Process 57
3. Carbon Based Processes 58
a. Reinluft Process 58
b. Sulfacld Process 61
c. Hitachi Process 63
4. Manganese Based Processes 65
a. Mitsubishi Process 65
b. TVA Osone-Manganese Processes 68
5. Selenium-Based Process 69
'a. Nor Deutsche Affinerie Process 69
b. Badische Anilin- and Soda-Fabrik 73
Process
6. Modified Chamber Process 75
7. Reversible Dry Absorbant Process 77
8. Process Cost Estimates 79
9. Comparative Evaluation of Flue 83
Gas Treatment Processes
References 105
Appendix I - Conversion Efficiency and Rate 115
Equation Graphs
Appendix II - Mathematical Models and 133
Computer Program Listings
Appendix III - Catalyst Data Sheets 155
Appendix IV - Capital Cost Sheets 219
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LIST OF FIGURES
Page
1 Computed Equilibrium Conversion Values 16
2 Results of Experiments of Boreskov, et al 28
3 Results of Eklund's Measurements 29
4 Results of Kinetic Measurements on Catalyst 3 30
5 Results of Kinetic Measurements on Catalyst 1 31
and 2
6 Monsanto-Penelec Process — New Plant 52
7 Monsanto-Penelec Process — Existing Plant . 53
8 Kiyoura-T.I.T. Process — New Plant 55
9 Kiyoura-T.I.T. Process — Existing Plant 56
10 Reinluft Process for S02 Removal 59
11 Lurgi Sulfacid Process for S02 Removal 62
12 Hitachi Process for S02 Removal 64
13 Flow Diagram of Mitsubishi Manganese Dioxide 66
Process
14 Flow Diagram of TVA Direct Sulfuric Acid 70
Process
15 TVA Direct Ammonium Sulfate Process 71
16 Nor Deutsche Affinerie Process for S02 Removal 72
17 Badische Anilin- and Soda-Fabrik Process for 74
S02 Removal
18 TYCO Modified Chamber Process 76
19 Gallery Dry Absorbant Process 73
20 Effect of Product Credit on Operating Cost of 87
Monsanto-Penelec Process (Existing Plant)
21 Effect of Product Credit on Operating Cost of 89
Monsanto-Penelec Process (New Plant)
22 Effect of Product Credit on Operating Cost of 91
Kiyoura-T.I.T. Process (Existing Plant)
23 Effect of Product Credit on Operating Cost of 93
Kiyoura-T.I.T. Process (New Plant)
24 Effect of Product Credit on Operating Cost of 95
Reinluft Process
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LIST OF FIGURES
Page
25 Effect of Product Credit on Operating Cost of 99
Mitsubishi Process
26 Effect of Product Credit on Operating Cost of 101
T.V.A. — Direct Acid Process
2? Effect of Product Credit on Operating Cost of 103
Gallery Process
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LIST OF TABLES
Page
1 Catalysis and Promoters for the Oxidation of 9
Sulfur Dioxide (Heterogeneous Catalysis)
2 Computed Equilibrium Conversion Values for 17
S02 + 1/2 02 tS03
3 Effect of Contact Time on Conversion 18
4 Forms Used for the Function f(PQ ,PSQ ,PSO ) 20
b £ j
5| List of Rate ;Equations Used in Calculations 21
6! Catalysts Used to Obtain Rate Equations in 24
Table 5
7 Rate Equations for the Oxidation of S02 Over 37
Platinum Catalyst
8' Methods for Radiation Enhancement 44
9 Effect of Radiation Pretreatment of Catalysts 48
for S02 Oxidation
10 Capital Requirements and Operating Costs for 80
S02 Oxidation Processes
11 Operating Cost Estimate Summary - Monsanto- 86
Penelec Process (Existing Plant)
12 Operating Cost Estimate Summary - Monsanto- 88
Penelec Process (New Plant)
13 Operating Cost Estimate Summary - Kiyoura- 90
T.I.T. (Existing Plant)
1^ Operating Cost Estimate Summary - Kiyoura- 92
T.I.T. (New Plant)
15 Operating Cost Estimate Summary - Reinluft 94
Process
16 Operating 'Cost Estimate Summary - Sulfacid 96
Process
17 Operating Cost Estimate Summary - Mitsubishi 98
Process
18 Operating Cost Estimate Summary - T.V.A. 100
Sulfuric Acid Process
1^ Operating Cost Estimate Summary - Gallery 102
Process
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I. SUMMARY OF LITERATURE REVIEW
The objective of Phase I of the program was to identify, by a
comprehensive literature search, and to evaluate existing and
potential methods of applying catalysis to the oxidation and
removal of the sulfur dioxide in powerhouse stack gas.
The literature notes a large number of individual materials, and
combinations of materials, as capable of converting sulfur dioxide
to sulfur trioxide. Unfortunately, the experimental conditions,
the conversion efficiencies, and the extrapolated economics for
essentially all cases cited were impractical in relation to the
present application. Among the numerous materials mentioned, the
better ones consistently were vanadia (V205), platinum, oxides of
iron, oxides of chromium, and carbon. Of these, the best, and
only practical ones, were shown to be vanadia and platinum, and
even platinum was shown to be economically impractical when com-
pared with vanadia. This conclusion is not surprising consider-
ing the depth and breadth of the original research that culminated
finally in the commercial vanadia catalysts now employed through-
out the world for sulfuric acid production. Nor is it surprising
considering the research effort expended during the past 30-odd
years to produce a catalyst better than vanadia.
For an equally long period, the kinetics of sulfur dioxide oxida-
tion and the catalytic mechanics have been studied by numerous re-
searchers.- Essentially all such studies were related to conversion
with vanadia catalyst in gas streams containing 5-12% sulfur dioxide,
typical of that in contact sulfuric acid plant operation. Seme half-
dozen kinetic equations are offered in the literature as descriptive
of sulfur dioxide conversion data with vanadia catalyst in relative-
ly narrow incremental ranges of conditions within the broad concen-
tration range noted above. Nothing of the sort was found relative
to conversion under conditions typical of the stack gas environment.
Considerable divergence of opinion was found with regard to the
mechanics of vanadia catalysis, particularly as to what constitutes
the limiting step. There appeared to be more support for the pro-
posal that re-oxidation of V*+ to V5+ is the limiting step than for
any other proposal.
Five of the kinetic equations reported for use with vanadia were
used to predict conversion efficiency at the concentration of sulfur
dioxide representative of that in flue gas. Subsequently, data
generated during the program with vanadia and simulated flue gas
showed two of the equations to be quite accurate in their predictions
for such conditions. These were an equation by Mars and Maessenvand
one by Eklund. The former, not having a factor to account: for equi-
librium, is very accurate in all regions below equilibrium and drifts
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off as equilibrium is approached. Eklund's equation includes a re-
sponse for equilibrium conditions and thus, while far less adequate
in sub-equilibrium regimes, becomes progressively more descriptive
as equilibrium is approached.
Based on the assumption that re-oxidation of vanadous oxide to vana-
dic oxide is the limiting reaction in the catalysis, a route is sug-
gested to low temperature vanadia catalysTir: It is known that the
activity of commercial vanadia catalyst begins at the temperature at
which the alkali pyrosulfate promoters begin to melt. It is also
known that the vanadia dissolves in the molten pyrosulfates and it
is assumed that such solubility is required for catalytic activity.
If this is true, a pyrosulfate system melting at a lower temperature
and dissolving vanadia should constitute an active catalyst system.
The alternatives are (1) the re-oxidation reaction of Vlt+ to V5+ is
extremely temperature dependent or (2) an entirely different mecha-
nism acts at lower temperatures. The literature did not provide
evidence to support or deny either of these alternatives.
The possibilities of enhancing catalytic activity through pretreat-
ment of either the feed gas or the catalyst by various types of ir-
radiation or by continuous irradiation of either were examined and
found to be impractical for two reasons: (1) beneficial effects
possible were transient and of low order; and (2) cost of treatment
was prohibitive, irrespective of the effect.
The recent literature notes several processes, proposed for the re-
moval of sulfur dioxide from powerplant stack gas ostensibly by
catalytic oxidation. On examination, not all of these processes,
at various stages of development, were found to involve catalysis
in the true sense of the meaning. Those employing vanadia, manga-
ous ion (In aqueous solution), and carbon were noted as being truly
catalytic, while others employing dry manganese dioxide or selenium
dioxide were noted as involving straightforward stoichiometric reac-
tions between sulfur dioxide and the "catalyst.11
All but two of the process types described in the open literature
were found to be impractical, or extremely limited in applicability,
on the basis of complexity (relating to capital and operating costs)
and "Product" value. Product value here refers either to market
value or to product disposal generally, i.e., substitution of one
pollution problem for another. The two process types of potentially
broad application are those employing vanadia and manganese dioxide.
These processes are represented by the Kiyoura and Monsanto Company
vanadia processes, producing ammonium sulfate and sulfuric acid, re-
spectively; and, the Mitsubishi manganese dioxide process producing
ammonium sulfate. As noted above, the latter process is non-
catalytic, which leaves only one type of potentially practical cata-
lytic process, namely, that employing vanadia catalyst.
2
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The Tyco process for sulfuric acid production from power plant
flue gases offers another possibility wherein S02 is converted
via a homogeneous catalysis route using, in part, the NOX con-
tained in the flue gas effluent. The process is still in a
developmental stage.
This review reflects only the processes reported in the litera-
ture prior to 31 December 1968. Subsequent developmental modi-
fications and new process schemes are necessarily not evaluated.
Future developments in the application of oxidative catalysis to
removal of sulfur dioxide are seen as evolving along the follow-
ing lines:
(1) Development of a "low" temperature catalyst.
(2) Improvement in economics and applicability through
engineering development.
(3) Inclusion of economical means of removing S02 from
flue gas to supply a more concentrated feed stream
to the catalytic converter.
Inclusion of economical means for simultaneously re-
moving the S02 and the NOX present in the flue gas.
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II- CONCLUSIONS FROM LITERATURE REVIEW
1. The materials consistently noted as effective In converting
sulfur dioxide to trioxide are: V205, Pt , Cr203, Fe203, and
L< o
2. Only V205j Pt, and C show sustained high conversion activity
below 500°C.
3. Although carbon appears to be active at temperatures below
400°C, no satisfactory method was noted for desorbing the
conversion product from carbon.
^ . No catalyst (other than carbon) was noted as having practi
cal activity below
5. The oxides of chromium and iron generally show sustained con-
version activity at 600°C and above „
6. The oxides of iron, chromium, and vanadium are essentially
"poison" resistant, while platinum is subject to poison-
ing by a variety of materials.
7. The activity of vanadia (V205) is maximized by promoting it
with alkali pyrosulfate, notably potassium pyrosulfate.
8. The reactions of vanadia catalysis are believed to occur in
a liquid (molten) phase of the pyrosulfate promoter in which
the vanadia dissolves.
9. The rate controlling step in vanadia catalysis is believed to
be re-oxidation of V+ ** to V+ 5 by oxygen.
10 o The kinetic equation by Mars and Maessen best fits the data
on non-equilibrium oxidation of sulfur dioxide over vanadia
catalyst under flue gas conditions; Eklunds kinetic equation
best fits the data at equilibrium.
11. The rate controlling step in platinum catalysts is believed
to be oxidation of sulfur dioxide by atomic oxygen.
12. The kinetic equation by Uyehara and Watson best fits the data
on non-equilibrium oxidation of sulfur dioxide over platinum
catalyst under flue gas conditions.
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13- Theories relating catalyst activity to pore volume, sur-
face area, lattice spacing, surface electronic effects
etc., are often erroneous.
14. Potential methods for enhancing catalyst activity through
irradiation by a variety of sources were shown to be im-
practical.
15- A vanadia based process appears to offer a practical means
for removing sulfur dioxide from powerhouse stack gas by
catalytic oxidation.
16. An adaptation of chamber process technology appears to offer
the possibility of simultaneous removal of SOX and NOX and
of retrofit to existing power plants (low temperature ex-
haust).
17. A process consisting of vanadia catalysis followed by a
sorption step to remove and concentrate S03 appears to
offer the possibility of improved economics compared to
the condensation route to S03 recovery.
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III. DISCUSSION
A. INTRODUCTION
This program was implemented by NAPCA with the objective of collect-
ing and evaluating all available data in the literature relating to
removal of sulfur dioxide from flue gases by existing or potential
methods that employ catalytic oxidation with subsequent recovery of
the sulfur value. There is a wealth of information in the literature
concerning the catalytic conversion of sulfur dioxide to trloxide,
but it is not directly related to the desired application.
Generally, the same materials that are effective as catalysts In
many applications, namely, the transition metals, are noted as active
in varying degree in effecting conversion of sulfur dioxide- There
is, inevitably, some combination of these and assisting materials
(promoters) that shows fair to good activity at some elevated temper-
ature. However, also inevitably, the activities are lower than, or
the temperatures are higher than, those associated with either
platinum or vanadium catalysts. The primary reason for such diver-
gence in behavior is connected with the stability of the sulfate of
the element employed as catalyst. To behave as an effective catalyst,
a material must release, or desorb, the conversion product, sulfur
trioxide. The transition metal oxides form sulfates with sulfur
trloxide and, except for vanadium, the decomposition temperatures of
the sulfates are usually above 500°C, frequently above 600°C, At
these temperatures, the conversion equilibrium begins to shift strong-
ly to the left so that, added to whatever inherent lack of conversion
activity these materials may have, is even poorer conversion inherent
in the reaction itself. The sulfate of vanadium, however, is un-
stable even at room temperature, and platinum does not form the sul-
fate under conversion conditions.
Catalysis generally is viewed more as an art than a science, and
this is reflected by two facets in the 'literature related to conver-
sion of sulfur dioxide. The first is the almost infinite number of
combinations of materials studied through the years for catalytic
activity. The second Is the almost amusing disagreement expressed
regarding the mechanism of the catalytic action. Whatever the vari-
ous views, however, essentially all of them relate to conversion in
gas streams where the sulfur dioxide concentration is generally In
the range of 5-12 vol % and not to powerhouse flue gas where the
concentration is In the range of 0.05 to 0.5 vol %.
When the literature refers to conversion under flue gas conditions,
it is describing one specific process among a number of processes
that have been proposed in recent years and pursued to. various
stages of development-. If a process was reported to produce a sul-
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• MONSANTO RESEARCH CORPORATION •
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fate, the process was examined in our survey. This revealed that
several of the processes were not truly catalytic, but since they
were examined anyway, they are reported along with those that are
truly catalytic.
Two NAPCA-sponsored processes are reported that were not in the
open literature. These are a) the Tyco process, which applies
Chamber Process Technology to flue gas cleaning, and b) the
Gallery Chemical Co. process which sorbs S02 from flue gas leav-
ing a contact converter and delivers it in a concentrated .stream
to an absorption tower. Both 'processes are interesting, but are
still in the laboratory stage.
Sulfur value is commonly .recovered as sulfuric acid or ammonium
sulfate. Overall economics appear to favor a heterogeneous cata-
lytic process recovering sulfuric acid in a concentration range of
77-83$ HaSOi*. While some other product may be Justified in particu-
lar instances, the potential supply of any product from this source
could affect the .disposal of the product. A vanadia contact pro-
cess, for example, recovering sulfur value is ammonium sulfate,
appears to afford attractive savings in capital cost. However,
in the United States at least the product is essentially valueless
and presents a disposal problem.
Finally, the practicality must be considered of integrating the sul-
fur dioxide removal process with the power production process. It .
is easy to visualize designing a new power plant to incorporate the
flue gas treatment plant. Integration of the treatment process with
existing power plant facilities, on the other hand, can only be
custon-engineered for each application. For this reason, costs of
such a treatment process for existing power stations can only be
given in rather bro.ad terms. In the future, another factor will af-
fect treatment cost; it is the requirement to remov NOX, or any
other pollutant, in the same treatment process for removing sulfur
dioxide. It must be recognized early that the .power industry is not
likely to be able to sustain a series of "black boxes" hooked on the
exhaust end of a power plant. Consequently, for any flue gas treat-
ment process, multifunctionality becomes an increasingly important
consideration.
In the remainder of .this report, the various aspects of the problems
touched on above are discussed in greater detail.
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B. CATALYSIS IN SO? OXIDATION
1. General Remarks
The reactions for producing sulfur trioxide and sulfuric acid have
been known for several centuries. One early reference, cited the
production and isolation of sulfur trioxide by Albertus Magnus in
the 13th Century (Ref. 1). At that time, however, sulfur trioxide
was made by the decomposition of sulfates. Phillips, in 1831
(Ref. 2), started a search for catalytic methods of oxidising sul-
fur dioxide to sulfur trioxide by observing that the reaction could
be effected by passing the reactants over "finely divided platinum,
heated to a strong yellow heat." These observations served as the
rudiments of the contact process of producing 803 through hetero-
geneous catalysis. Homogeneous catalysis for converting S02 to SO 3
is about a century older. The rudiments of the chamber process
were known as early as 17^0 (Ref. 3).
Despite the fact that these reactions were used both in the labora-
tory and in the commercial production of sulfuric acid for several
hundred years, the importance of catalysis in the economics of these
processes only began to be studied in detail in the twentieth
Century. Therefore, the literature review for this program was
limited to information published since 1900 on the .catalytic oxida-
tion of sulfur dioxide. Even though the community at large for
several years has expressed concern about the air pollution problems
resulting from the burning of fossil fuels, comparatively little ap-
pears in the literature concerning the application of technology to
remove pollutants from flue gas. This being the case, all references
in the literature dealing with the catalytic oxidation of sulfur di-
oxide to sulfur trioxide were reviewed.
The catalysts and promoters that seem most promising for gas-solid
heterogeneous catalysis of S02 oxidation, on the basis of literature
references, are listed in Table 1. Rhenium metal has also been re-
ported as a possible catalyst for the oxidation of sulfur dioxide,
but, until recently, has been too expensive to be economically feas-
ible. Recent increased availability has forced the price of rhenium
to the level.of economic feasibility and its catalytic potential
should be reevaluated (Ref. ^). Generally, the catalysts listed in
Table are not effective below 400°C, although Mandelik (Ref. 5)
claims 98$ conversion in two stages "substantially below iJ000C" with
a vanadium pentoxide catalyst promoted with potash and phosphorus
pentoxide. Bienstock (Ref. 6) reported tests with potash-promoted
vanadium pentoxide in which 98$ conversion of sulfur dioxide in flue
gas (between 300° and 370°C) was claimed. However, from what has
been seen to date, there appears to be no heterogeneous, gas-solid
catalyst system which operates at temperatures below 300°C.
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Table 1
CATALYSTS AND PROMOTERS FOR THE OXIDATION
OF SULFUR DIOXIDE (HETEROGENEOUS CATALYSIS)
Catalyst Promoters
V205 Bi203; alkali metal sulfates, oxides, and
pyrosulfates; Ag2SOt+> P205
Pt Ag, Au, Ni, Pd
Iron, Iron oxides CuO
Cr203 Sn02, A1203, Ti02, BaO
Carbon
Solid-liquid catalysis in sulfuric acid solution is a second type of
heterogeneous catalysis operating at lower temperature. The S02 is
absorbed in dilute sulfuric acid and catalytically oxidized while
still in solution. Two catalysts are currently being used for this
type of catalysis: selenium dioxide (Ref. 7), and manganese salts
such as MnSO^ (Ref. 8). In the course of the oxidation, the catalyst
is reduced and must be reoxidized in a separate step. It might be
argued that these processes are not truly catalytic, but they will
be discussed herein for the sake of completeness.
All the catalytic processes that are homogeneous, gas-phase reactions
are variations of the chamber process for the production of sulfuric
acid, with nitrogen oxides acting as the oxidizing agents. Other
types of catalysis which can operate at low temperatures are photo-
oxidation, electrolysis, and oxidation in a spark discharge. Al-
though Buff and Hoffman (Ref. 9) had oxidized S02 in a spark dis-
charge as early as i860, no commercially feasible methods employing
high energy, light, or electric currents for catalyzing this reaction
have actually been put into practice. A related type of homogeneous
catalysis is effected by bombardment of the reactants with high
energy particles such as photons, neutrons, x-rays, etc. These
oxidative processes will also be considered in this report.
2. Factors Affecting Activity of Solid Catalysts
Since heterogeneous catalysis must of necessity involve reactions
which occur on the surface of the catalyst, surface area is important
in determining its activity. However, there are only certain portions
of the total surface area of a catalyst which are able to adsorb the
reactants and allow the reaction to occur. In other words, the total
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surface of the catalyst contains areas of high catalytic activity,
i.e., active sites. Two catalysts with equal surface areas, there-
fore, may have different activities resulting from a difference in
the concentration of active sites on their surfaces. A related pro-
perty of catalyst particles which affects activity is porosity.
The outer surface of a catalyst is the first surface poisoned, but
it makes up only a fraction of the total surface area of a porous
catalyst pellet. Therefore, it is often the inner surface area of
the pore walls that determines catalyst efficiency. For extremely
fast reactions, the diffusion of the reactants through the pores
of the catalyst is the rate-determining step.
Another set of physical properties also thought to affect activity
results from defects in crystal lattices which, in turn, tend to
result in active sites on the catalyst .surface. Several investi-
gators have reported evidence of effects on catalyst properties of
lattice imperfections such as dislocations and point defects
(Ref. 10 through 14). Since there is a direct relationship between
point defects in crystal lattice and the semiconductor properties
of a crystal, it has been suggested that the semiconductor proper-
ties of catalysts could be used to predict their relative activi-
ties (Ref. 15). A good correlation, if one could actually be
found, between semiconduction and catalytic activity would be a
powerful tool in catalyst design. Thus far, no such correlation
has been reported. The relationship of catalyst activity, to cry-
stal defects is intimately connected to the question of the effect
of radiation on catalyst activity. The activation of catalysts by
irradiation is universally felt to result from the creation (by
the radiant energy) of lattice defects in catalyst crystals. How-
ever, on .prolonged exposure to elevated temperatures, these defects
tend to be "annealed" out with loss of the added activity.
Over thirty years ago Eyring (Ref. 16) and Horiuti (Ref. 17) cal-
culated the magnitude of the activation energies of adsorption of
hydrogen on various substates. They concluded from their studies
that the lattice spacing would play a dominant role in determining
the magnitude of the activation energy of adsorption. This work
led to the concept that lattice spacing and the spatial arrangement
of atoms in a catalyst crystal would affect the activity of the
catalyst in a particular reaction. This is the basis of the so-
called "geometric factor" as a parameter in determining catalytic
activities. These effects are taken to result from the fact that
the geometry of the species adsorbed on the catalyst surface fits
with varying degrees of ease into the templates formed by the ar-
rangement of atoms on that surface. The results of the simple
case treated by Eyring and Horiuti, could be explained on the basis
of lattice spacing, but the activities of more complicated catalyst
systems are not easily predicted on the basis of the "geometric
factor."
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3. Comparison of Platinum and Vanadium as Catalysts
During the period of the development of vanadium catalysts in the
United States (1927-30), attention was given to the following ob-
jections to the platinum catalysts then in use: (a) they were ex-
pensive, (b) they were extremely sensitive to poisons even in
minute amounts, and (c) their conversion efficiencies were low.
These same reasons are often given today to explain why the use of
platinum catalysts has been completely supplanted in the contact
process over the last three decades. However, we felt that it
would be worthwhile to review these arguments and see if technolog-
ical advances during that;period would not have some effect on the
validity of the arguments brought forth in support of the use of
vanadium masses in this process. This seemed particularly
necessary since we could find no published evaluation of the re-
lative merits of these catalysts since the outbreak of World War II,
A number of papers were published in the early thirties comparing
the catalytic properties of vanadium and platinum masses (Ref. 18-
23). The state of the arguments as of 1936 was fairly accurately
summarized by Far lie (Ref. 2^4) as follows:
I. Two Viewpoints of Comparison.
The two groups of catalysts may be compared from two dif-
ferent standpoints with regard to cost of catalyst, name-
ly, (a) on the basis of manufracturej and (b) on the
basis of manufacturing cost plus royalties.
II. Disadvantages Attributed to Vanadium.
1. Thompson (Ref. 18) making the comparison from standpoint
"b", brought out that the prices charged for vanadium
have been much too high, compared with the prices of
platinum masses.
2. Thompson also emphasized the disadvantage of vanadium in
that it handles only 7 to 8 percent S02 gas, as compared
with 10 percent gas handled by platinum. Vanadium cata-
lyst can handle 9 or 10 percent S02 gas, as has been dem-
onstrated with Calco mass and on a laboratory scale with
Selden mass; but in plant practice, the large majority of
vanadium-mass plants actually do operate on gas within
the range 6 to 8 percent S02.
3. Neumann (Ref. 19) and Streicher (Ref. 20) have pointed
out the disadvantage of vanadium, as compared with plati-
num, in respect to overloading capacity.
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4. Vanadium and platinum masses are affected unfavorably by
accumulations of dust within the catalyst bed.
5. When worn out, vanadium mass has no salvage value.
III. Advantages Claimed for Vanadium
1. Vanadium maintains its conversion efficiency undiminished
for a longer period than platinum.
2. The average conversion efficiency under normal loading is
higher for vanadium than for platinum.
3- Vanadium catalysts are immune to "poisoning" by arsenic,
chlorine, and some other elements which harmfully affect
the activity of platinum masses. However, for the reasons
set forth above, such immunity to poisoning has become of
relatively slight importance.
<*. On the basis of manufacture cost, vanadium is initially
cheaper per daily ton of ^SOt* than platinum. However,
when the salvage value of platinum is taken into account,
this advantage is materially reduced. Also, with high
royalty charges added to the cost of manufacturing vana-
dium mass, the advantage of lower initial cost of cata-
lyst is yielded.
5. Vanadium mass offers less operative trouble and anxiety
than platinum, especially when operating with metallurgi-
cal gases.
6. The latest vanadium masses are hard, rugged, and able to
withstand handling without crumbling.
7- By means of vanadium mass, any country possessing supplies
of vanadium among its natural resources is independent of
foreign countries for raw material for the manufacture of
sulfurlc acid catalysts. This is a decided advantage in
time of war.
IV. Disadvantages Attributed to Platinum
1. Platinum, even if not "poisoned", suffers a gradual decline
in activity with use over a period of years, and at least
(within 10 years) especially in the primary converter.
2. In a primary converter the platinum catalyst life is like-
ly to be shorter than that of vanadium mass.
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3. Platinized asbestos and magnesium; sulfate are subject to
"poisoning" by arsenic, chlorine, etc.
4. Platinized asbestos is fragile and liable to be damaged
in handling or to be compacted in use.
5. As mentioned, both groups of catalysts become choked by
deposits of dust within the bed.
V. Advantages Claimed for Platinum
1. Platinum can be reclaimed from a spent mass, to the ex-
tent of 90 percent of the weight originally introduced.
2. The price of platinum has fallen to 25 percent or less
of its 1920 price.
3. The quantity of platinum in the modern platinized cata-
lysts has been reduced to 2 or 3 troy ounces per daily
ton of ^SOti produced.
4. The manufacturing cost of modern platinum catalysts is
in the neighborhood of $150 to $200 per daily ton of
H2SOit, including the cost of the platinum, of which over
90 percent is later reclaimed.
5. Platinum catalysts have a much .higher capacity for over-
loading at sustained conversion efficiency than vanadium
masses.
6. In plant practice, platinum handles gas containing from
2 to 3 percent more SC>2 than the vanadium masses most
widely employed. This means substantially larger pro-
duction capacity and smaller power expense, for a given
plant investment, with platinum catalysts.
7. Platinized silica gel is reported to be immune to arsenic
poisoning.
8. According to U.S. Patent 1,384,566 with hot gas-purifi-
cation, other platinum catalysts are not poisoned by
either arsenic or sulfuric acid mist.
On the basis of the above arguments, Thompson, who published the most
extensive article on the subject, concluded that platinum was the
better of the two catalysts.
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Since World War II, a great deal of progress has been made in tech-
nology. Therefore, it is worthwhile to reevaluate the platinum ver-
sus vanadium controversy in the light of these advances. First,
the early platinum catalysts had surface areas of only several
square meters per gram; today, platinum catalysts are produced with
surface areas of 200 square meters per gram.
Further, vanadium catalysts in 1939 were being installed in only
new plants while it was generally only the older plants which used
platinum catalysts. It may be that some of the differences in the
production figures obtained from these two types of installations
resulted from better plant design rather than better catalyst mass.
The problem of the largest concentration of S02 that can be handled
economically in the feed stream by the two types of masses might be
an important consideration in the design of a manufacturing plant
for H^Oi,. However, in oxidizing flue gases of fixed lower level
S02 concentration, this criterion is less meaningful.
4. Conversion of SOg Under Equilibrium Conditions
The degree of conversion of S02 to S03 under equilibrium conditions
can be calculated by the procedure outlined by Dixon and Longfleld
(Ref. 25). These authors assume, however, that the S02 feed gas is
produced by burning sulfur in air at two atmospheres pressure. In
flue gas the equations representing the partial pressures of the
components of Interest are as follows:
[b-0.5a(l-y)]Pt
02
l-O.Sa(l-y)
P a-yPt
rS02 '
so
l-O.Sa(l-y) (2)
a(l-y)Pt
-
l-O.Sa(l-y) (3)
where: Pt = total pressure of all flue gas components (atm.)
Px = partial pressure of component x in flue gas (atm.)
a = initial fraction of S02 in flue gas
b = initial fraction of 02 in flue gas
y = fraction of SOg remaining unconverted at equilibrium
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The equilibrium constant of the reaction
S02 + 1/2 02 = S03
is given by equation 4.
KP
Pso2
If we combine equations (l)thru (^) we obtain-.
0.5a[Pt.Kp2 -l]y3+[Pt-Kp2(b-0.5a)-(l-1.5a)]y2
+[2-1.5a]y+[-l+0.5a] = t)
(5)
The cubic equation (5) can be solved for y and for a range of values
of a, b, Pt and Kp. Only values of y between zero and one will have
physical meaning. These y values can then be substituted into equa-
tions 1, 2, and 3 to give values of PQ,, PsOs • and Pso3« Tne frac-
tion of S02 converted is of course 1 - y. Tne values of Kp used in
solving these equations were calculated from the following equation
given by Dixon and Longfield (Ref. 25).
K = "'956 _ 4.678
Kp RT R (6)
where: T = temperature (°K)
R = gas constant
The calculated equilibrium conversion versus temperature is plotted
in Figure 1 for a flue gas containing 0.3 mole percent S02 and
several different concentrations of 02.
The equilibrium conversion values start to drop below 100$ for tem-
peratures of ^750°K C477°C) or higher. At each temperature, the
presence of more initial oxygen results in higher values of the equi-
librium conversion. Similar computer calculations were made for dif-
ferent initial concentrations of S02. For example, at 2i5 vol
of 02, initial, the equilibrium conversion values were computed au
650°, 800°, and 900°K, respectively, and are shown in Table- 2.
Details of the calculations are shown in Appendix I.
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a
B
n M
r
100
,u 90
S 80
HE
z 70
O
£ 60
IU
Z 50-
O
i40
2 30
a 20
2
10
0
600
CONVERSION OF SO2 TO SO3
02= 10.0%
O2= 2.5%
xXV o,
**. V ^^^ °2
•N^>r-
•%xV^Os.
= 7.5%
02=5.0%
650 700
750 800 850 900 950 1000 1050 1100
TEMPERATURE °K
Figure 1. Computed Equilibrium Conversion Values
-------
Table 2
COMPUTED EQUILIBRIUM CONVERSION VALUES FOR: S02 + 1/2 02*S03
S02 Temperature, °K
Initial Vol. % 650800900'
0.1 0.993 0.811 0.525
0.2 0.994 0.843 0.523
0.3 0.994 0.842 0.521
0.4 0.995 0.842 0.518
0.5 0.995 0.842 0.516
0.6 0.996 0.842 0,513
Below 800°K, the changes in equilibrium conversion values are insig-
nificant in view 01' the accuracy inherent to the computer evaluation.
At 900°K, the equilibrium conversion value drops with increasing
initial concentration of S02 in the flue gas feed. This agrees with
a statement in (Ref. 26):
"Increasing the S02 content of the feed results in a lower
equilibrium oxygen concentration and a lower equilibrium
conversion of S02 to S03 at any given temperature."
It is realized that these calculated equilibrium conversion factors
are of limited value for two reasons: (1) the desirable temperature
range for operation is below 900°K, a temperature range at which the
equilibrium conversion factor is close to 100/S, and (2) the short
residence times in a practical reactor for flue gas treatment pre-
clude the attainment of equilibrium conditions.
5. Conversion of S02 Under Non-Equilibrium Conditions
a. Introduction
The primary objective in this phase of the search was to obtain
kinetic data on the catalytic oxidation of S02 under flue gas condi-
tions, that is, very low concentrations of S02 and 02, compared to
the usual concentrations of these gases in gas mixtures processed
with the contact method for producing sulfuric acid. It was realized,
even before this program started, that not much literature was avail-
able pertinent to the primary objective. A secondary objective, then,
was to assess the literature on catalytic oxidation under contact
process conditions (8-10$ S02, 20% 02). This assessment would consist
of a critical study of the kinetic models and kinetic data presented in
various papers on the subject.
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To achieve this goal, a wide variety of kinetic equations appearing
in the literature (Ref. 27-H) were analyzed and are discussed in
some detail in subsequent sections in this volume.
b. Oxidation over Vanadium Catalysts
Napier et al., (Ref. 27) discusses the results of a preliminary study
of the catalytic oxidation of S02 under flue gas conditions. The
basic composition (by volume) of the gas mixture was:
C02 = 13.5$, 02 = il.7%, N2 = 76.3%, H20 = 5.IX and
S02 = 0.1%
An industrial vanadium catalyst was used. The catalyst contained
V205 and K2SOl4 with a silica support and ranged in composition from
6.5-7-5 weight percent V20b and 6.3-10.6 weight percent K20. The
authors found that one of these catalysts (6.7% V205, 9.5% K20) yield-
ed a fractional conversion of 93% at 429°C and of 82% at 511°C.
Conversion was found to be independent of contact time, in the range
studied, as shown in Table
Table 3
EFFECT OF CONTACT TIME ON CONVERSION
Weight of Contact Fractional
Catalyst, g Time, Sec. Conversion., %
3.1 0.43 9^
2.0 0.27 9^
1.4 0.19 98
0.6 0.09 94
A limited amount of work at other concentrations of S02 and 02 showed
that varying the S02 concentration over the range 0.3^/5 - 0.27% (by
volume) did not produce any large changes in the fractional conversion
The fractional conversion was consistently over 90%. The variation
of conversion with oxygen content was studied at two additional
oxygen concentrations: 6.8 vol % and 1.85 vol. % 02. The fractional
conversions obtained were equal for all practical purposes, .i.e.,
90% or better.
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The main value of this exploratory work is, to quote Napier,
"firstly, to confirm that the presence of water vapor
and carbon dioxide in the gases has no adverse effect
on the catalyst, and secondly to show that the required
contact time at low sulphur dioxide concentration is
much lower than that used in the contact process" ......
and (it) suggests that it might be technically feasible
to oxidize sulphur dioxide in boiler flue gases."
In the main approach to the problem of kinetics of oxidation of S02
over V2C>5 and of obtaining a rate equation, two general methods have
been followed. First is the empirical method. This method involves
assuming that the rate is given by
r = k f (PSo
2>
and then by fitting the rate with assumed forms of the function
f(pso2> po2>
one obtains a rate equation. Various forms have been chosen for the
function f as shown in Table ^ . These rate laws usually hold for
certain conversion ranges and temperatures. Most were not purported
to hold through wide ranges of conversion of S02 to S03. Other equa-
tions were reported that were supposed to hold through wide ranges of
conversion. Some of these are shown in Table 5,'
and were used in calculations described later. Tne catalysts used to
obtain the kinetic data which were fit to these equations are listed
in Table 6.
The second method employed the application of the "Theories Method"
of Hougen and Watson (Ref. 45). These theories assume that gas-solid
catalytic reactions take place by the following steps:
1. Diffusion of the gaseous reactants in the gas phase to the
catalyst surface.
2. Adsorption of the reactants upon the surface.
3. Reaction between the adsorbed reactants on the surface.
4. Desorption of the products.
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Table 4
FORMS USED FOR THE FUNCTION f(Pn Pon , POA )
Author and Reference
A. Boreskov (36)
B. Boreskov ( 37)
C. Boreskov ( 38)
D. Calderbank (3*0
E. Zakerevski & Chang (
F. Davidson-Thodos (^3)
Krlchevskaya
H. Boreskov
p p
S02 - S02
SO
SO.
'SO:
P f S°2
-------
Table 5
LIST OF RATE EQUATIONS USED IN CALCULATIONS*
A. P. H.. Calderbank
ln
i/;
12.07
so
1 -
/2 v
Applicable temperature range 370° - 500°C
B. P. B. Eklund
,1/2
C,
SO-
V
k P,
S0
1-f
lpso
Y
PS03
P l f z K
./
values of k obtained from table of Eklund:
r(°C)
k(x!0b)
500° 460 420 483
27-0 10,3
2.0 19.0
Applicable temperature range 420-550t'C
Mars and Maessen
V
kKpso2 Po2 I —
2/3 x 10
-8
exp
0
RT
All rates above are in moles S02 per second per gram of catalyst
and all pressures are in atmospheres. In all equations the value
of Kp, the equilibrium constant for the gaseous reaction S02 + l/2
•*SO-j is given by
In K.
22.600 21.36
RT ~ R
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Table 5 cont'd
for catalyst 1:
, . -26.000 . 15.89
I n \f = * 4- —— —
in K RT R
for catalyst 3:
, ,
ln
-21.000
= RT
10.89
R
Applicable temperature range:
D. Davidson-Thodos**
lJOO-500°C
V =
ln
in
p P !/2
rS02 02
SO
P l/2 K
*
/K
RT
in
in
RT
in /K— = 57,750 .
n ^n ni1
U 2 nl
In K
SO
3
RT
R
39.^2
R
R
R
in K - 36,350 . 29.77
ln KN2 Rf^ + ^R^
Applicable Temperature range:
**The factor 1/3600 is used to convert the rate as given to moles
302 per second per gram. The temperature here is in °Rankine
and R is in Btu/degree-mole.
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Table 5 cont'd
E. Goldman
—
In k
in KQ
In Ks
in Ks
k P P 1/2
S02 S02
_ 15,709
RT
_ 28,769
2 RT
02 RT
03 RT
Pso3 \
1 ' '
\p P 1/2 K
. r o r\ * r\ *** t
S02 02 p 1
/
h KS02 PS02 + KS03 PS03)
15.72
R
27.82
R
20.58
R
5*1.82
R
Applicable temperature range:
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2
O
z
Z
-I
o
n
m
c/>
m ro
> -t-
3
o
I
o
O
3
TJ
O
H
O
z
Table 6
CATALYSTS USED TO OBTAIN RATE EQUATIONS IN TABLK 5
Equation
Cat. Size
P. H. Calderbank 10-14 mesh
R. B. Eklund
Mars & Maessen
Catalyst #1
Catalyst #3
Davidson-Thodos
Goldman
0.67-1.38 mm
0. 5 mm
0. 5- mm
3/16 in. spheres
60-80 mesh
Catalyst Composition
V205
11.3%
K
0
Other
H.6% Na
Prepared by method given in
Maxted, E.B., "Catalysis
and Its Industrial Applica-
tions," p. 24? (London;
Churchill) (1933)
Commercial catalyst - no
analysis given
Q.2% V205 5.1% Ca;
175? K 0.9% Na
The (K+Na)/V (atom/atom) is 2.8
6.8% V205
10* K
0.7? Ca
2.9% Na
The (K+Na)/V atom/atom) is 2.5
•v8* V205 10% K20 0.5% Fe
Commercial catalyst -
Davison Chemical Co., Code 902
-------
Assuming that one step is rate controlling, Hougen and Watson derived
rate equations for various type reactions. Various workers have fit
S02 oxidation rate data to the equation that assumes the surface re-
action is rate controlling. Equations D and E in Table 5 were arrived
at by this method. In one of these (E), it is assumed tuat N2 in the
stream has no effect; in another (D), the nitrogen is included. The
reason for considering N2 is that it may compete with S02 and 02 for
adsorption sites. The term in brackets in D and E in Table 5 supposed-
ly takes account of the reverse reaction of 863 decomposition. This
term also appears in other equations that are arrived at by a more
empirical approach than the Hougen-Watson approach.
It is obvious from Tables ^ and 5 that no accord seems to exist in
the explanation of the rates of catalytic S02 oxidation for I^SOi,
production (Ref. 27, 31, and HJ|).
The most recently published rate-law equation for S02 oxidation is
that of Mars and Maessen (Ref. 28 and 29). They have shown that it
is possible to fit the data of Eklund (Ref. 31) and of Boreskov
(Ref. M) to their kinetic equation (see Equation C, Table 5). This
is a step in the right direction.
The basic assumption of Mars and Maessen is that, in the catalytic
system, the following equilibrium is established continuously and in-
stantaneously :
SO? + 2V+5 + CT2 Z S03 + 2V+1* (7)
This reaction is assumed to be a two phase equilibrium between the
gas phase and the "solid" phase. The V* 5 , V*1* and 0~2 are believed
to be dissolved in a thin liquid film on the surface of the catalyst.
This liquid film is believed to be composed of melted alkali or
alkaline earth pyrosulf ates . Tandy (Ref. 47) has presented evidence
that this melt actually does exist for alkali-promoted catalysts in
the temperature range ( iJOO°-500°C) studied by Mars and Maessen.
The equation above is assumed to have an equilibrium constant:
PS03
K =
J7" Pso2ao-'
where a = activity coefficient
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It is assumed that the activity coefficients of V+1* and V+5 are equal
and that a.Q-2 is constant so that it may be incorporated into K to
give the equilibrium constant:
[V+"] p
SO 3
K =
tv+53 PSo2
(9)
The rate of reaction is controlled by the rate of oxidation of V"1"1*
to V1"5 by 02 in the gas phase as follows:
02 -f v+l* = V+5 + 02~
and the rate equation of the forward reaction is given by:
V = k Pn [V+"]
°2
m
(10)
Using (9) and (10), it is possible to show that the rate is:
= k P,
/KP
S0<
1 m
The best fit of the rate data indicates that the value of m is near
2 so that m = 2 is chosen. Equation (11) is the same then as
Equation (C), Table
The contention of Mars and Maessen that equilibrium [Eq. (7)] is
rapidly established may be true, but, from their discussion of their
experiments, it seems that one can only say that equilibrium is
established but not how rapidly it is established^ The question of
how rapidly it is established may not be very important, since a
steady state condition may explain the situation Just as well.
The assumption that equilibrium exists is reasonable because the
equilibrium constant can be directly determined and calculated fro'til
kinetic data to obtain fair agreement. Further the fact that their
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equation fits the data of two other investigators lends support to
their treatment.
Kadlec and Regner (Ref. 147-1^9) report that their kinetic data sup-
port the assumption of Mars and Mae.sson that the oxidation of V+l+ is
the rate-determining step, but they picture this process as happening
in the three steps given below:
(a) 02 +
V+5 + 02~
(b) 02~ + V+1+ ? V+5 + 2 0-
(c) 0- + V+l+ t V+5 + O-2
They feel that, consistent with their data, (b) is the rate-controlling
step.
Of Equation (C) in Table 5 is algebraically manipulated, it can be
written in the form.
/P -,1/2
0-
V
~J
1/2 J/2/P \
-.-^ *(-) (—)
• k \kK ' Pe '
1/2
(12)
A graph of (P0?/V)1/2versus (pS03/pso?>1/2 should yield a straieht
line.
The results of this operation for Mars and Maessen's data, Eklund's
data, and Boreskov's data are shown in Figures 2-5 taken from
Reference 29. These graphs indicate that Equation (12) pre-
dicts the data reasonably well for higher temperature but that
some deviations are observed at lower temperatures. How low a
temperature is needed seems to depend upon the particular cata-
lyst used. Boreskov's data are at ^85°C only, so temperature
comparison is not valid in this case. It should be noted that the
lines obtained by plotting data according to Equation (12) are
almost parallel for different temperatures in the high temperature
range. This indicates that the expression (1/kK)1/2 is almost inde-
pendent of temperature. This is seen to be approximately true from
the expressions for K and k of Equation (C) in Table 5- At lower
temperatures, this term changes with temperature and the lines at
lower temperatures have different slopes than for higher temperatures
They also have a curvature that appears to depend on the
ratio.
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J?0.
180.
140
100
60. _
V'".1"
A
-1C -12 -OB -04 O 04 00 1.2 16 20 2.4
Figure 2. Results of Experiments of
Boreskov, et al
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12ft
100
Q7S
7
050
1^37*
029
.18 -16 .14 .12 -10 -08 -06 -04 -02 0 02 04 00 OB 10 12 1
I 10 18
Figure 3- Results of Eklund's Measurements
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SO -IS -1.0 -05 0 OS 10 15 20 2%
Figure 4. Results of Kinetic Measurements on
Catalyst 3
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900.
800.
/5
600
500
400.
300
900
100
c* »•
• AM'C I Cololyltl
• <»0t J
• 4
-------
The Mars and Maessen equation will not be valid when one approaches
equilibrium for the reaction S02 + 1/2 02 * S03 because no provision
is made for the rate to approach zero as equilibrium is approached.
Their graphs, however, indicate validity of their equation up to
about 85% conversion.
Several investigators (Ref. 28, 44) have reported that if one plots
the logarithm of the specific raue constant, or of the rate itself,
versus reciprocal temperature, two straight lines of different slope
occur in different temperature ranges. The break in the Arrhenius
plot occurs between 400°C and 500°C and seems to depend on the com-
position of the gas phase and the catalyst used.
The break in the curve indicates a probable change in mechanism.
This could be explained by the fact that between 400-500°C, the
pyrosulfate melt forms and that the mechanism of catalysis is differ-
ent above the melting point than below the melting point. This melt-
ing point may be affected by the amount of V+l* (as VOS04) present
which is determined by the gas composition according to Equation (7).
Mars and Maessen, in a 1964 publication (Ref. 28), seemed to agree
with the existence of a break in the curve, but, in their latest
paper (Ref. 29), they are not as sure and indicate that their rate
equation and data do not predict a break in the Arrhenius plot.
However, the fact that some deviation from their equation at lower
temperatures exists would indicate that the possibility of a break
in the Arrhenius graph is high. Their data is consistent with the
change in mechanism occurring by a phase change on the outer surface
of the catalyst. For a solid outer surface, thp equilibrium in
Equation (7) above may not exist or the assumptions leading to
Equation (9) may not be valid and so the derivation of Mars and
lessen equation may not be valid. The break in the Arrhenius graph
is further supported by data from Calderbank (Ref. 34) who studied
the adsorption equilibrium of 02 and S02 on a K2S04 promoted V205
catalyst. He found that between 400°-500°C the type of active centers
occupied by the adsorbed molecules on the catalyst surface changed.
The point of change depended, for a given temperature, on the pressure
of 02 or S02 (whichever was being adsorbed). A change in the type of
adsorption centers could cause a change in mechanism and consequently
a change in rate. Calderbank found that 02 (AHa(js = 6,4 kcal/mole)
is much less strongly chemisorbed than S02 (AHads = 28.8 kcal/mole)..
He also studied the rates of adsorption of S02 and 02 on the catalyst
surface but, due to their magnitude,did this at lower temperatures
(250-300°C) than those at which the equilibrium adsorption and S02
oxidation are done (400-500°C). From these, he concludes that the
mechanism for oxidation is: .
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(a) S02 + 2e(surface)-»-S02 — (fast chemlsorption) (13)
(b) S02 — + 02 •* S03 + 0 (slow second-order reaction (14)
between gaseous 02 and ad-
sorbed S02)
(c) 0~~ + 1/2 02 + 2e(surface)(fast desorption) (15)
(d) SO 3 (surf ace J-^SO 3 (gaseous ) (fast desorption) (36)
He chose reaction (14) as the slow reaction so that the rate is:
V = k'P~ aS02
°2 (17)
and fits the data best to:
aS02 = k" P<,n O'"
S°2 (18)
so that
V = k P " •** P
?so2 Po2 {J9)
where k = k 'k"
This equation holds only for low conversions (<15#). In a later
paper (Ref. ^8;, Calderbank expands Equation (19) to fit cases for
high percent conversions by taking account of the decomposition of
SO 3. He did this by choosing the exponent of PS02 as 0.5 above in
stead of 0.4 (which he Justified by saying it is sufficient for de-
sign purposes). He then wrote the rate as!
v Vkl Pso2 Po2 - k* f (pso2' V Pso3} (20)
and chose the function f so that V = 0 at equilibrium and
k]/k2 = Kp. This gave Equation (A) in Table
Of the remaining equations in Table used In the calculation de-
scribed below, two of them (D and Ey are derived from Hougen-
Watson theory assuming that the rate of reaction between surface ad-
sorbed S02 and 02 is the rate-controlling step. The main difference
in these two equations is that one theorizes that N2 competes for the
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active site (D, Davidson-Thodos) while the other does not (E,
Goldman, et al). Since Calderbank (Ref. 3*4) found that N2
is not adsorbed by the catalyst he used at the temperature
above 400°C, it seems that Equation (E) in Table is basically
more correct than Equation (D). This is in addition to the fact
that Equation (D) predicts a decrease in rate with increase in
temperature, an unlikely situation. The investigation by Goldman,
et al (Ref. 30) also studied the relative effects of fluidized
and fixed beds on the rate. They fitted both sets of data to the
same rate equation but found that the rate was 3-1* times higher
with fluidized beds than for fixed beds. Both Goldman, et al,
and Davidson and Thodos used total pressures of over one atmos-
phere; these did not exceed 1.5 stmosphere for Davidson-Thodos
and 2,5 atmosphere for Goldman, et al. Since all rate equations
are in partial pressures of reactants, this is probably not criti-
cal for these small differences. Equation (B) in Table was pro-
posed by Eklund (Ref. 3D and was arrived at semi-empirically.
The Hougen-Watson theory influenced his derivation but was modified
to fit the experimental data. He found that the value of the
specific rate constant may vary slightly with the catalyst size.
The value used is for commercial-sized catalyst pellets, which
are cylinders approximately 1/3 inch in diameter by 1 inch long.
The form (H) in Table ** is used by Boreskov (Ref. 36, HH) and
is supposeu to be valid over wide conversion ranges.
A general theme of most of the papers on oxidation kinetics is
that the data is not presented in its entirety in table form, thus
preventing the reader from using it if desired. Usually graphs
of the data are drawn but, in the case of Boreskov and -others,
the graphs are not large enough for reading and are incompletely
labeled.
The present state of catalytic oxidation of S02 seems to be the
following:
V205 catalysts promoted with alkali metal compounds are
the most popular and active catalysts, K being the most
commonly used promoter for economic reasons. The reaction
is believed to take place, at least partly, in a liquid
phase of pyrosulfates on the surface in which the V205
dissolves. The V+5 is at least partly reduced to V+l4 in
the reaction, and the relative amounts appear to depend
on the relative amounts of S03 and S02 in the gas con-
tacting the catalyst, the temperature, and which and
how much promoter is used.
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There appears to be a change in mechanisms of catalysts between
iiOO-500°C which probably is related to the phase changes of the
pyrosulfate promoter. The relationships between these phenomena
are not known and most work has been done for S02 concentrations
between ^% and 15$ by volume.
The rate equations in Table 5 are purported to be valid for high
conversions and the relevant constants were available. These
rate equations were used to calculate conversion of S02 as a
function of time and also as a function of W/F, where W is
the weight of catalyst required for a given conversion and P
is the flow rate in moles of S02 per second. The calculations
were performed with a computer and are discussed in Appendix
II.
The calculations were made at flue gas conditions
(Pn = 0.028 atm, Pcri = 0.003 atm and PM = atm
U2 o(J2 N 2
initially)
for the purpose of comparing the various equations in regard to
their predictions about contact time and W/F values. Calculations
were made at temperatures from 375° to 500°C. None of the equa-
tions were valid over the whole temperature range; however, in
some cases the equation was assumed to be valid outside the re-
ported range by 25°C or so.
In AppendixII, Figures 1 through 7 show the percent conversion
of S02 versus contact time at various temperatures as calculated
using the various rate equations. Figures 8 through 1*1 show
the percent conversion of S02 versus W/F.
It is seen in the graphs that the two catalysts for Mars and
Maessen and Eklund's data give curves that group more closely
together. This is expected since Mars and Maessen fitted
Eklund's data to their equation. If calculations could have
been made with Boreskov's equation, it should also be near
the values given by Mars and Maessen, and Eklund's equations.
c. Oxidation Over Platinum Catalyst
A list of the equations that have been used historically to fit
kinetic data for S02 oxidation over a platinum catalyst are
given in Table 7.
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Lewis and Ries (Ref. 49) tried to fit Equations (A) and (B)
in Table 7 to experimental data that they generated on the
rate of oxidation of sulfur dioxide over platinum under a
variety of conditions. They reported three types of experi-
ments. In the first, S02 was oxidized in air. Both the
temperature and concentration were varied in these experiments,
the temperature from 400°C to ^50°C, and the initial concen-
tration of S02 from 0.109 to 0.519 mole percent. .
In the second type of experiment, the sulfur dioxide was present
at about the same concentration as in the first type, but the
concentration of oxygen was reduced by dilution with nitrogen.
The temperature was maintained at 450°C.
The third type of experiment employed gas mixtures consisting
of 0.5 mole percent or less of S02, 1 to 5 mole percent of
SO3, and the remainder oxygen and nitrogen.
Lewis and Ries concluded that kinetic Equation (A) in Table 7
fit their data and the data reported earlier by Knletsch (Ref.
50). They also concluded that their Equation (A) could be made
to fit the data more accurately than could Bodenstein's (Ref.
51) or an equation derived from the law of mass action. Un-
fortunately, no attempt was made by Lewis and Ries to deter-
mine the reaction constant and thus the rate of reaction in
some sort of standard units, such as moles converted per second
per gram of catalyst. In fact, during their experiments, they
did not determine the weight of catalyst in the reactor.
Roiter and co-workers (Ref. 52) investigated the oxidation of
S02 over platinum catalyst under equilibrium and non-equilibrium
conditions by use of radioactive S^502. They were able to
fit their data to Equation (A), and, on the basis of the fit,
proposed two alternative mechanisms for this reaction:
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Table 7
RATE EQUATIONS FOR THE OXIDATION OF S02 OVER PLATINUM CATALYST
A, Lewis and Ries (Ref.^9)
B.
-dm kP
dt
S02 [In
-In
SO.
SO-
M = moles of S02
Px = partial pressure of component x (atm)
k = constant
Superscript 'e' denotes equilibrium conditions
Bodenstein (Ref.51)
PS°2
dt" = k 71
c = concentration of S02 (moles/cc)
C. Roiter, et al. (Ref.52 )
P,
p . S p
- \, rn_ rQH-_
dt
p . 5 p
02 rS02 -
so.
Ke = equilibrium constant
£>. Uyehara and Watson (Ref .53 )
V =
O U
)
p . 5 p
02 S02
SO 3
e
K
R = Gas constant (cal/mole-°K)
V = reaction rate (moles S03 produced/sec - g of catalyst)
k = 2.77 x 10~4 C-8000/T + 14.151*)
}2 = e(20,360/RT-23.0/R)
S03 = (16.800/RT-17.51/R)
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I II
(a) 02 - 2 0 (a) S02 + 02 * S03 4- 0
(b) 0 + S02 + S03 (b) 0 + S02 * S03
(c) 2 0 + 02
In each case, reaction (b) was considered as the rate-determining
step. They were unable to distinguish between the two mechanisms,
using their data. To do this they would have had to make a detailed
investigation of the relationship between reaction rate and oxygen
concentration.
All of the experiments in this study by Roiter, et al were performed
with the concentrations of sulfur dioxide and oxygen appropriate
for a contact sulfuric acid plant. When their expressions are applied
to the oxidation of S02 under flue gas conditions, the contact times
predicted to obtain 90% conversion of the S02 are an order of magnitude
too high compared to laboratory test data generated on this program.
A modification of the reaction constants is afforded by the equation
of Uyehara and Watson (Ref.53), (see equations C and D in Table 7).
The k of the Roiter equation is replaced by a more complicated ex-
pression that varies not only with temperature but also with the
chemical composition of the reaction mixture. Consequently, when a
calculation is made of the percent conversion of S02 as a function
of contact time, the results are more in keeping with contact times
and conversions actually realized in practice.
d. Oxidation of SO? on Chromium Oxide Catalyst
Another substance which historically has been known to be effective
in catalyzing the oxidation of S02 is Cr2C>3 in conjunction with other
metal oxides. Rienacker (Ref.5^) noted that, although neither Cr203
nor Sn02 alone had any appreciable activity in this regard, their
combination was a good contact catalyst for the oxidation of S02?
High conversions of sulfur dioxide were reported through the years
by several investigators (Ref.55-59)for mixtures of Cr203 with such
metal oxides as SnO,, Ti02, and Fe203. The additons of BaO and
Fe20-< with a mixed Cr203-Sn02 catalyst were found to increase its
efficiency, and the addition of CuO, MgO, SrO, ZnO, Bi203, Mn0.2, CoO
and CuO were found to inhibit its activity (Ref.56), NiO and A120H
seem to vary in their effect on a Cr203 catalyst, depending on the
other metal oxides present (Ref. 55-57).
These catalysts are reported to have typical activation energies of
from 19 to ^5 kcal/mole between 400-500°C which is comparable to
energies of activation obtained with platinum and vanadium catalysts.
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Kurin and co-workers (Ref. 60, 6l) have studied the kinetics of the
oxidation of S02 on a Cr203-Sn02 catalyst and have fit their results
to the following equation:
o r\ o A ~cr\
b03 S02 b02
= k (.^i I
dt
where: C^ is the concentration of component x
GX is the concentration of component x
if equilibrium conditions had been obtained
k is the rate constant of the reaction
They measured the rate of reaction at various concentrations of Sn02
and Cr203 and also at various temperatures. No support material was
used. They found that there were maxima in the catalyst activity
which occured at between 2 and 3% Sn02 and between 30-50? Sn02. The
plotted values of k versus 1/T indicate a break in the curve at
about 440°C. These authors attributed this break to a change in
mechanism at that temperature. They postulate that the mechanism of
the reaction taking place below i|^0°C is given by reactions a-c and,
above this temperature, by reactions d-f, below:
(a) Cr203 + Sn02 t 2Cr02 + SnO
(b) 2Cr02' + S02 * Cr203 + S03
(c) 2SnO + 02 +• 2Sn02
(d) 2(Cr203 2S03) Z 2Cr203 • 3S03 + S03
(e) 2Cr203-3S03 + Sn02 * 2(Cr203 • 2S03) + SnO
(f) 2SnO + 02 t 2Sn02
There is little in these literature citations that would allow us to
decide a priori whether Cr203 based catalysts could compete with V205
catalysts for use in flue gas oxidation processes. In a later section,
the poison resistance of these catalysts will be discussed showing
V205 and Cr203 catalysts fairly equivalent in this respect.
e. Oxidation over Iron Oxide Catalysts
A fourth type of catalyst, iron oxide, has, historically,been used al-
most exclusivply by the Soviet Bloc countries for the oxidation of
S02 (Ref. 62-77).
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The reaction can be carried out in twc ways to oxidize S02. At low
temperatures 350°-^50°C, (Ref. 73 and 76) sulfur dioxide is oxidized
over Fe203 to form iron sulfate. This latter material is then
passed through a furnace where hot gases at 750°-850°C decompose the
sulfate and entrain the S03. The S03 is then absorbed in sul-
furic acid.
At higher temperatures 600°-700°C,(Ref.64-66,69)a more truly catalytic
process is effected in which S02 is converted to S03 as it passes
through a Fe20s catalyst bed,sometimes with a CuO promoter. It has
been shown, however, that iron sulfate is often an intermediate in
this reaction, although it is produced and decomposed in a single
process step.
Boreskov (Ref.66) has studied the kinetics of the reaction and found
by x-ray examination that above 670°C the catalyst shows only the
Fe203 spacings. Below 670°C, it contains up to 415? of S03. The sul-
fated catalyst is probably a mixture of Fe203 and Fe2(SOl4)3. At 680°C,
Boreskov found that the rate of reaction was given by:
f
1>5
S02 \ , , x
1 p 1 1
PS03 ' °2 k2 )
/PS03
• p
• so2
0 .5
/
V = k, U . Pn fr-^-U • (22)
Another variation of the catalytic oxidation over iron oxide process
is one in which the S02 is oxidized in an iron sulfate solution (Ref.68)
The reaction consists of two steps. In the first FeSO^ is oxidized
to Fe2(SOit)3. In the second Fe2(SOi,)3 is reduced to FeSOi, with
the formation of H2SOi4. During the early stages of the reaction, the
optimum temperature is 60°-80°C, and the optimum ratio of S02 tp 02 is
1-1-5. For the second step, the nntimum temperature is 80°-90°C, and
the optimum ratio of S02 to 02 is 1:4.
6. Catalyst Poisoning and Effects of Inert Gases
Some discussion of the relative resistance to poisoning of the
platinum and vanadium catalysts is found in this report in the
section on the comparison of these catalysts. Further review of the
effects of inert gases and poisons on the catalyst efficiencies of
these systems is provided by Donovan (Ref.78).
The molecular weight of the inert components such as helium, C02 and
N2 has a significant effect upon the oxidation rate of sulfur dioxide
over vanadium catalyst; the heavier inert gases giving higher rates.
This effect has been attributed to the transfer of kinetic energy
from inert gas molecules striking the catalyst.
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The great number of studies on the effect of water vapor on the ac-
tivity of vanadium oxide as. well as platinum catalysts have produced
conflicting results. The only definite conclusion which could be
drawn from these studies is that water vapor is certainly not harm-
ful and probably is beneficial to the catalytic oxidation of sulfur
dioxide over metal oxide catalysts, at least as long as temperatures
are kept high enough to prevent water and sulfuric acid from condens-
ing on the catalyst (Ref. 78 and 79).
The poisoning of catalysts by foreign substances occurs by two widely
differing mechanisms. Poisoning occurs most commonly as a result of
physical coating1 or choking of the catalyst and the bed voids by
foreign dust. Small amounts of iron sulfate and sulfuric acid in the
gas stream tend to aggravate the situation by cementing the catalyst
and dust together into an almost impervious mass. This necessitates
periodic removal of the layer of catalyst nearest the entrance to the
reactor and replacement of it with fresh catalyst.
Another type of poisoning results from chemical reactions between
components in the gas stream and the catalyst surfaces. The most
effective of these catalyst poisons, as far as the S02 oxidation
reaction is concerned, is arsenic and its compounds. Susceptibility
of platinum catalysts to arsenic poisoning was one. of the reasons
cited for switching to vanadium catalyst in the contact process. It
is reported that 80,000 times more arsenic is needed to deactivate
vanadium catalyst, than to deactivate platinum catalyst. Chromium
oxide catalysts are reported to have roughly the same resistance to
arsenic poisoning as vanadium (Ref.80) and can be regenerated by
treatment with carbon monoxide.
The halogens and their compounds are another important class of
catalyst poisons. Both chlorine and hydrogen chloride have been
found to seriously impair the activity of platinum catalysts. Near-
ly full activity can be restored by passing heated air over the
catalyst, but some platinum is volatilized in the presence of chlorine
necessitating periodic addition of fresh catalyst.
The halogens, especially chlorine and fluorine, are felt to be in-
jurious to vanadium catalyst if present in large amounts and for long
periods of time. This loss in activity is thought by many to be due
to the volatilization of vanadium or V205 in the presence of halogens.
In contact plants, platinum catalyst masses have lasted as long as 10
years or more before regeneration was required. There have been in-
stances, however, where such a mass only lasted for 2 1/2 years.
Vanadium catalyst, on the other hand, is replaced at the rate of 1/16
of the mass per year, the equivalent of 17 years of catalyst life.
Vanadium catalyst is normally not regenerated.
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No accurate estimate of the life of a Cr203 mass can be stated because
this type of catalyst is not used in commercial contact acid plants.
Little is reported on the stability of Fe2C>3 to the attack of poisons.
However, it is known that the catalyst gradually.degenerates because
of the formation of iron sulfates (Ref. 8l).
7. Oxidation of S02 in an Electrical Discharge
Ever sJnce it was shown by Buff and Hoffmann (Ref. 82) and von Wilde
(Ref. 83) that sulfur dioxide and oxygen could be passed through an
electrical discharge to form sulfur trioxide, this reaction was stud-
ied in the hope that it would have some commercial importance. A
good review of the work in this area prior to 1939 is that by
Glockler and Lind 'Ref. 84}} and a more up-to-date review is that of
Gregory (Ref. 85)
Mahant (Ref. 86) found that S02 could be oxidized to S03 in the
presence of oxygen using an electrodeless discharge. He obtained the
maximum conversion (30-40%) with a mixture containing 60% S02 and 40%
02. Mahant reports that no thionates or polythionates were formed by
his process. It is clear that it would be impractical for any com-
mercial process to operate with such a high S02-to-02 ratio and with
such low conversion yields.
Poliakoff (Ref. 8?) demonstrated that if oxygen alone is subjected
to the action of a discharge, it will unite with sulfur dioxide after
being removed from the influence of the discharge. Sulfur dioxide
was not activated in the same way. This led to the supposition that
the reaction proceeded through the attack of ozone molecules, known
to form in a discharge, on the S02 molecule.
Nechaeva (Ref. 88) observed that the yield of S03 formed in a high
frequency discharge decreased linearly with an increase of S02 con-
tent and increase in gas velocity. He found similar results (Ref. 89)
for a high voltage arc discharge. However, for practical purposes,
he reports that the highest obtainable yield is 66-70% by his method.
In a later paper, Nechaeva (Ref. 90) investigated the effect of
the wavelength of the radiation on the oxidation of S02 by air in a
high frequency discharge. He found that, by changing the wavelength
of the radiation, he could change -the products of the reaction. At
wavelengths of 320 m, 03 (and then presumably S03) were formed
primarily, while at a wavelength of 236 m the major products are
nitrogen oxides.
All of the above were laboratory scale investigations of the oxi-
dation of S02 in an electric discharge. Browne and Stone (Ref. 91)
studied this reaction from the point of view of its possible commer-
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cial use for removing S02 from flue gas. Their conclusion was that
0.09 kw-hr/ft3 of gas are required to obtain 99% conversion of S02
in flue gas. A 75%. conversion would require 0.0*1 kw-hr/ft8 of flue
gas. Based on the 991 conversion figures above, the power consumption
and conversion cost per ton of S03 produced are 280,000 kw-hr and
$1700 per ton, respectively.
These calculations were made using the assumption that the flue gas
contains 0.3$ S02 and that power costs $0,006 per kw-hr. This cost
is astronomical compared to the present cost of H2SO»t($32/ton, 66°
Baume; which is approximately equivalent to an 803 cost of $40/ton).
Browne and Stone's report, that they got appreciable conversion of
S02 in their test apparatus even with the corona discharge shut off
casts some suspicion on their data.
However, even using the more optimistic figures of Gregory (Ref. 85)
who reports power expenditures for a variety of materials produced
in a corona discharge (not including S03) of 8000-25,000 kw-hr/tona
we obtain costs for SOa production of $48-$150 per ton.
The conclusion that may be drawn from these studies is that, although
it is possible to produce 803 from S0£ and air in an electric dis-
charge, the fact that H2SO«t can be produced cheaper by other means
makes this method impractical on a commercial scale,
8. Feasibility of Photochemical & Radiation Methods for S03
OxidatiolT" ~~
a. Introduction
This section deals with the technical and economic feasibilities of
using photochemical and radiation methods to promote or enhance the
oxidation of S02. A general lack of pertinent data on these methods
as applied to S02-oxidation makes it difficult to accurately assess
technical and economic feasibility. On the basis of some general
theoretical and engineering considerations, augmented by the limited
experimental data in the literature, some conclusions can be drawn
regarding the practicality of these methods. It seems appropriate
to list here the various types of applicable radiation and the methods
in which they can be used. (See Table 8 below). Practicality or
impractical!ty of each will be shown in the following sections.
b. Direct Irradiation of S02-Containing Flue Oases
Untreated or pretreated flue gases can be exposed to radiation such
as (a) ultraviolet, (b) particles, (c) particles and electrons,
(d) rays, and (e) neutrons.
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Table 8
Methods for Radiation Enhancement
Irradiation Pre- Irradiation of
Direct Irradiation Treatment of S02 Containing
Type of of the S02- Catalysts for Oases in Presence
Radiation Containing Gases S02-0xidation of Catalysts
Ultra Violet X
Alpha X XX
Beta X XX
Gamma X XX
Neutrons X X X
One method of evaluating the feasibility of using direct radiation
is to estimate the power or activity requirements for this method
of S02 conversion. For a limiting case, which will yield a low esti-
mate of the power or activity required, the following assumptions
can be made:
• all absorbed energy is used uniquely in promoting S02
oxidation, i.e., there are no transmission losses, no
radiation leaks through the vessel wall, and no radiation
energy wasted in producing chemical changes other than
the one desired.
• the yield, i.e., the number of chemical changes produced,
is proportional to the total radiation dose.
• the yield depends only upon the total dose absorbed
rather than the radiation intensity. It will also
be assumed that the intensity, i.e., the particle or
photon energies, is such that ionization of species
does take place.
The proportion of molecules changed can be estimated by:
P • 10~6'G-M'R (Ref. 92) (23)
where
P = proportion of the molecules changed
G = "0" value, i.e., the number of chemical bonds affected
by radiation per 100 electron volt absorbed by the
.specimen
M = molecular weight of the Irradiated species
R = dose rate in megarads
• MONSANTO RESEARCH CORPORATION •
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If all of the S02 must be converted, I.e., if P = 1, then the dose
rate (h) required is :
Excluding irradiation by ultraviolet photons for the time being
and limiting the discussion to the other forms of radiation listed
above, the needed dose rate can be evaluated if. the G value of the
S02 oxidation reaction is known. No data were found in the litera-
ture on G values of this reaction. However, for many chemical reac-
tion? in which the total chemical change is proportional to the
radiation do.se absorbed, a G value in the order of 3 is common.
For example, the energy required to produce one ion-pair in air
IL= about 3^-5 eV, and the corresponding G value is thus close to
3- It is likely that G will be approximately 3 also for our ca::e
since the oxidation of SC>2 proceeds subsequent to the production
of oxygen ions. If we take M as 64, then
R - " 5,200 megarads,
a very high dose indeed. The strength of the radiation source-
necessary can be estimated through the following considerations and
assumptions:
- 1 kilowatt hour of energy corresponds to 3.6 x 1013 ergs.
- A radiation dose of 1 Mrad corresponds to an energy absorp-
tion of 108 ergs per gram of irradiated material.
• - Therefore one kilowatt hour fully absorbed can treat 79*J
pounds per Mrad, or x Ibs for 79^ Ib-Mrad/Kw-hr.
- In general, the .following relationship exists
x = 79R'W (25)
where
x = lb of product irradiated per hr with dose of R Mrad
R = dose in Mrad
VJ = kW output of radiation from the source.
• MONSANTO RESEARCH CORPORATION •
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Assuming a power plant emitting 1MMSCFM of flue gas containing
0.5% (volume) of S02, the quantity of S02 to be oxidized is 106gm
moles of S02 per second or 5^,000 Ib of S02 per hour. If, in
keeping with previous assumptions, it is assumed that only the
S02 ultimately benefits from all radiation, W, from equation (25),
is calculated to be 350,000 kw. A radiation source rated at
350,000 kw output must be provided to handle the oxidation of S02
from a power plant which releases 1MMSCFM of flue gas with 0.5?
(volume) of S02. If the radioisotope source were to be a brems-
strahlung producer such as Sr90, 6.5 x 107 kilocuries of it would
be required at an estimated future cost of about $13. 2 x 109
(Ref. 93). Also, the quantities required would simply not be
available; the projected total productions of Sr90 is approximately
300 kw by 1975 (Ref. 93). If the isotope were used as SrTi03 its
weight would be k.^ x 106 Ib,
From these tentative considerations it can be concluded that direct
irradiation of flue gases by ionizing radiation is not technically
or economically practical. The picture is even less attractive if
such factors as radiation losses, shielding and handling problems,
depth of radiation-penetration problems, etc. are taken into account
Literature on photochemical oxidation of S02 using ultraviolet
(UV) photons is also scant but not nonexistent as was the case
for ionizing radiation. In Refc 9^ are examples of such literature.
However, measurements of the reaction parameters were typically
performed at conditions markedly different from those prevailing
in flue gases and the- literature data thus may not be pertinent.
According to Ref. 9^, quantum yields for S02-02 mixtures varying
in ratio0of concentration from 1:2 to 2:1 and at UV wavelengths
of i860 A and 2070 X, vary from about 0.5 to 3.1. The highest
yield occurs at the most energetic wavelength, i.e., at i860 &.
If such a yield could be achieved in flue gases containing small
quantities of S02, 2.1 x 1025 ultraviolet photons (A=l860 A) per
second would be necessary to handle the 0.5 vol % of S02 in a
1MMSCFM power plant. This corresponds to an energy of 2.25 x 101U
ergs or 2,25 x 10U kw. It should be stressed, however, that'this
calculation is based upon data valid for an S02-02 mixture. It
is likely that in flue gases the photons will also interact with
the other much more plentiful species and the energy requirement
of the ultraviolet source could be very much higher than calcu-
lated here. The photochemical approach to S02 oxidation does
not appear to be attractive. Very large ultraviolet sources would
be necessary and their cost would be prohibitive even if the units
are available at all. Another problem might be the reliability
of the ultraviolet sources at flue gas temperatures.
'46
• MONSANTO RESEARCH CORPORATION •
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c. Irradiation Pre-Treatment of Catalysts for S0?-0xldatlon
Reference 95 is an excellent survey of the effects of ionizing radia-
tion on catalysts. Most of the catalyst systems discussed are not
pertinent to the problem of S02 oxidation,. Table 9 , extracted from
Ref. 95, summarizes research on radiation-treated catalysts for S02
oxidation. The platinum catalysts on asbestos substrates show an
increase in activity subsequent to exposure tc x-ray doses of 1-2 x
1018 eV per gram of catalyst. The Increase in activity is shown
to disappear after a relatively short period of time (12-24 hr).
Operation of the catalyst at temperature apparently gradually re-
moves the effects induced by the ionizing radiation. It is evident
that the effects of a radiation treatment are of no use in a flue
gas reactor, which must operate for extended periods of time. Two
examples of radiation-treated V205 catalysts are listed in Table 9.
There appears to be some indication that the activity of V205 cata-
lysts on a diatomaceous earth substrate Is slightly increased if
operated at temperatures below i!00°C. Whether or not this slight
Increase in activity exists for extended periods of time could not
be ascertained since the literature source cited in Ref. 95 could
not be obtained. The nuclear reactor-irradiated V205 catalyst
showed a decrease of activity as compared to the untreated samples.
It is perhaps appropriate to stress again the high cost of radiation
treatment. Even if the activity of a V205 catalyst could be improved
10% by a neutron dose rate of M x 10;7 nvt, such improvement would
be prohibitively expensive and not practically realizable. For
example, neutron fluxes in a thermal reactor irradiation chamber
are typically on the order of 10Jl* neutrrns'cm -sec. Thus, irra-
diation periods of 4 x 103 seconds would be necessary to achieve
an integrated flux of 4 x 10l' nvt. For a reference power station
it was calculated that 20,000 cu ft of V?05 catalyst would be
necessary- If a 10% improvement could be achieved, 18,000 cu ft
of radiation-treated catalyst would be required, Total irradiation
chamber space, in nuclear reactors in the USA, is usually in the
order of a few cu ft or less,. Irradiation charges are typically
on the order of $7 per cu in,per hr. The cost of the treatment is
thus prohibitive, even if the nuclear reactor time could be made
available.
From these considerations, it appears that radiation pre-
treatment of catalysts for S02 oxidation is not very promising
Many other possible disadvantages of radiation pretreating could
be cited. For example, in neutron pretreatment a slight amount
of radioactivity can be induced which, in a large quantity of
catalyst, could result in major shielding and health problems.
• MONSANTO RESEARCH CORPORATION •
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Legend: X:
Y :
n :
51/35:
X-Rays
Gamma Rays
Neutrons
Changed from 35 to
t : Increase
4- : Decrease
51
Table 9
Tesla Disch: High Voltage Disc
RT: Room Temperature
a : Day
• EFFECT. OF RADIATION PRE-TREATMENT OF
2
o
•z.
(ft
•z.
H
o
m
rn -cr
> CO
o
i
o
TO
o
H
O
Z
•
Catalyst
Pt-
asbestos
Pt-
asbes tos
V205-K20
(A) on
Diat .earth
(B) on Si02
gel
V2C3
V205-
Cr203
(0-1)
V205_
Ag2SOl«
-Si02
Reaction Radiation
S02+02, X18
260° 2xl018
e v | gm i n
moist air
S02+02, X100
300° IxlO18
e v | gm in
moist air
S02+02 Y3xl020
400-550° evjgm
Tesla disch,
H2 + D2 y^xlO20
-78° ev|gm-78°
CH3OH+02 nl.2X1020nvt
-••HCHO thermal
S02*air n4X1017nvt
460-520°
CATALYSTS FOR S02
Results E
Yieldt ,51 j 35, Return
to normal in 24 hr
Yieldt ,94 | 88. Return
to normal in^!2 hr
(A)Slighttat 400°,
slight+above 400°
. (B)-x-205Uat 400°,
slight+above 400°
t ( 5-10 ) | 1
No effect on pure or
doped V205. Doped
sample 30-50$ more
active than pure
*21*,520°;*33*,
460°
OXIDATION
xplanation
Radiolysis
products of
H20
-
-
-
Effect below
detection if
proportional
to Cr
Transmuta-
tion* or
"polishing by
Remarks
No increase in dry air
No increase in dry air
Tesla disch. also effec-
tive for ( A ) , not
stated for (B)
Stablold, RT
0.03%Cr by transmuta-
tion
# 1 1 OHl^cr
n
-------
9. Irradiation of S02 Containing Flue Oases in the Presence of
Catalysts
A limited amount of literature is available on catalysis In the
presence of ionizing radiation. However, no literature was found
on catalytic oxidation of SC>2 in the presence of radiation. It
would seem that most of the considerations with respect to high
cost of large radiation sources and of awkward handling and health
hazards discussed above would be equally valid here
10. Liquid Phase Catalysis
The oxides of iron and manganese are both utilized catalytically
in aqueous solution to effect the removal of 80$ from power plant
flue gas. However, in each case, the process appears to be unworthy
of further development at this time, due to high capital and operat-
ing costs. In addition, each represents a high degree of complexity
in operation, relative to other proposed processes.
11. Homogeneous Gas Phase Catalysis
Literature reporting the use of gas phase catalytic oxidation of S02
was concerned with various applications of the nitrogen oxides. These
schemes suffered commonly from high capital and operating costs as
presented. Problems include increased residence time (i.e. large
gas handling equipment) and loss of catalyst out the stack, which
creates additional pollution problems.
• MONSANTO RESEARCH CORPORATION •
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C. APPLIED CATALYSIS IN FLUE GAS TREATMENT
1. General
A review of the various processes associated with waste gas clean-
ing made it necessary to examine any process from which the product
was a sulfate. The applications of catalytic oxidation to flue gas
treatment, as noted in the literature, fit into three groups on the
basis of catalyst type: (1) vanadia, (2) carbon, (3) manganese. These
three materials appear to meet the criteria of a catalyst in this
application. However, in the case of manganese, only manganous ion
is truly a catalyst. Where manganese dioxide is employed in some
processes, its action is not truly catalytic because one of the
defining criteria for catalysis is that the material remain unchanged
chemically at the end of the reaction. It appears more likely that
the reaction between sulfur dioxide and dry manganese dioxide pro-
ceeds as follows:
4 + 4 + 2 + 6+
S02 + Mn02 »• MnSOu (26)
rather than
Mn02 + 1/2 02 + S02 - Mn02 + S03 •• MnSO^ + 1/2 02
(27)
Reaction (26) is a straightforward, stoichiometric, redox reac-
tion in which manganese is reduced and sulfur is oxidized.
It became expedient to define catalytic oxidation processes, for the
purposes of this review, as those in which sulfur dioxide (or sulfurous
acid) is oxidized to sulfur trioxide (or sulfuric acid) in the presence
of a material which remains chemically unchanged at the end of the re-
action.
All of the processes will be presented with the understanding that
only those involving the vanadia, carbon and manganous ion are truly
catalytic processes.
A homogeneous catalytic process, a modification of the chamber process
is also described although it is not yet in the open literature.
The cost estimates presented for the various processes are our in-
dependent estimates based on published information only, standard
estimating procedures, and the NAPCA guidelines. We did not use
estimates found in the literature for one size plant and scale
these estimates up to a 1^00 megawatt station.
50
• MONSANTO RESEARCH CORPORATION •
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For some of the processes, there was Just not enough information
to permit educated guessing on the basis of sound engineering
judgment. This is particularly evident in relation to the wet
carbon and TVA processes. For one version in each of these cate-
gories, there was Just enough data to tempt us to "guesstimate"
and the cost estimates show it. Consequently, we estimated only
one version in each category.
2. Vanadia Based Processes
a. Monsanto-Penelec Process (Ref. 96, 97, 98, 99, 100)
The Monsanto-Penelec process consists of catalytic oxidation of
flue-gas sulfur dioxide to sulfur trioxide, which is recovered as
sulfuric acid. In 1961, the Pennsylvania Electric Co., Monsanto
Company, Air Preheater Co., and Research Cottrell Corp., constructed
and operated a small pilot plant at Penelec's generating station
at Seward, Pa. Following this, a prototype plant was sponsored by
Metropolitan Edison and Monsanto at a generating station in Portland,
Pennsylvania. The Monsanto-Penelec process described here, though
similar, is not identical with the Monsanto-Metropolitan Edison pro-
cess, because details of the latter process are proprietary.
The Monsanto-Penelec process is essentially a contact sulfuric 'acid
plant, modified to give high heat economy and low-pressure drop.
For high temperature effluent applications (Figure 6) the flue gas
is taken from the boiler at 850°-900°F and passed through a high-
temperature electrostatic precipitator where virtually all (99.0%)
of the fly ash is removed. .The gas then flows through a bed of
vanadium pentoxide catalyst, with a residence time of 0.3 sec, where
the S02 is oxidized to S03. The gas is then cooled to about 200°F
by stepwise passage through an economizer and air preheater. Cooling
causes the formation of sulfuric acid mist by reaction of the sulfur
trioxide with moisture in the flue gas. Since the condensation be-
gins in the air preheater, both it and the mist eliminator, used to
remove the remaining acid, must be made of corrosion-resistant materi-
als .
Sulfuric acid will not condense and cause_corrosion if the tempera-
ture is above approximately 40.0°F. It has been, recommended that
corrosion-resistant materials be used in all areas where the flue
gas is below 500°F at design throughputs. The 100°F allowance is
provided because the metal will be cooler than the gas being cooled
and also the flue gas may be at a lower temperature when operating
at conditions other than full load.
51
• MONSANTO RESEARCH CORPORATION •
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2
O
z
en
>
Z
O
m
(A
o
o
o
a
TJ
O
X
O
Z
ro
HOT STEAM TO
AIR TURBINE
FUEL.
BOILER FEED WATER
FLUE GAS
8SO—9OO°F
S02
ELECTROSTATIC
PRECIPfTATOR
o
Figure 6. Monsanto-Penelec Process — New Plant
-------
STEAM
TOTUIMME
•OUR
ftSD WATER
o
z
01
>
z
-i
o
a
m
o
i
o
o
X
s
H
O
HOT AMI-
FUEL
MMLE*
PNEOMTATOM
Flue Gas Race: 2.5x10* SCFM
Gas flOM. vol.
Stream
H,
C02
HjO
Oj
SO,
S0j
"°,
CCFH10n
i
71.90
14.70
7.25
2.30
0.30
-
0.05
150.0
2
7«.95
10.70
7.25
2.75
Trace
0.30
0.05
ldo.3
3
75.28
11.83
6.98
2.36
Trace
Trace
0.05
It.?. 9
Figure 7. Monsanto-Penelec Process -- Existing Plant
-------
Figure 7 is a diagram of the Monsanto-Penelec process for an exist-
ing power plant as envisioned from various descriptions in the
literature. The material balance would be the same for both high
and low temperature effluent applications. The obvious point of
departure and the major problem in applying the process to existing
stations is the requirement of preheating the feed to the converter.
We have assumed here that steam will be used as the heating medium.
Additional capital would be required for an oil burner to provide
hot gas for the heating medium. The possibility of using hot flue
gas from the primary boiler for this purpose remains to be explored
in detail. The latter would require judicial heat swapping, assum-
ing the appropriate mechanical modifications could be made. In any
event, it appears that application of the process to low temperature
effluent may be more costly per installed Kw in capital and operating
costs than application to high temperature effluent.
The process has several outstanding advantages, i.e., no moving parts
are required, problems of recycling an adsorbant are obviatedj the
final temperature of the flue gas remains essentially unchanged (a
problem of considerable magnitude in wet scrubbing processes), and
the SOa is easily recovered as sulfuric acid.
There are also some disadvantages. The air preheater and mist elimi-
nator must be made of corrosion-resistant material. The saleability
of the relatively dilute, 705? to 80% sulfuric acid may present prob-
lems, although it may be feasible to concentrate the acid. The
electrostatic precipitator for the high temperature effluent from
the boiler must be large since it is in the hot zone preceding the
converter. Also, the process appears to be more easily applied to
construction of new plants than to modification of existing power
plants.
b. Kiyoura - T. I. T. Process (Ref. 97, 100, 101, 102)
A process similar to that of Monsanto-Penelec is currently under
development oy Dr. R. Kiyoura at the Tokyo Institute of Technology
(T.I.T.). The process consists of catalytic oxidation of S02 by
vanadium pentoxide catalyst, but instead of producing'eulfurid acid,
sulfur trioxide from the converter is reacted with gaseous ammonia
to produce solid ammonium sulfate. The Kiyoura - TjIiT, proceed is
presently in the pilot plant stage of development* A large pilot
plant, capable.of handling 500-1000 cu m/hr of flue gae, was don-
structed in 1967 and tied into a process steam boiler at the Toyo
Koatsu fertilizer plant in Omuta(Kyushu), Japan.
For high temperature effluent applications (Figure 8)j flue gas
moves from the boiler at about 850°F to a dust collector and then
to a catalytic converter where S02 is oxidized to S03 in a fixed>-
bed of vanadium pentoxide catalyst. After cooling to about 500°F,
the gas stream receives ammonia which prevents condensation of sul-
fUric acid, since it reacts to form ammonium sulfate.
• MONSANTO RESEARCH CORPORATION •
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STACK
NH,
HOT AIR
2
O
z
v>
Z
H
O
D
ni
v_n
a
o
i
O
O
X
T)
O
O
Z
FLUE CAS
88O°F
O.3% SO,
AIR
PREHEATER
OR
ECONOMIZER
72O-85O°F
BAG FILTER OR
ELECTROSTATIC
PRECIPJTATOR
DUST CATCHER
O 0
AMMONIUM SULFATE
Figure 8. Klyoura-T.I.T. Process — New Plant
-------
STEAM TO TURBINE
O
Z
(/>
>
Z
•H
O
a
fl
JO
o
i
o
o
a
•o
o
O
Z
8OLER FEED WATER
STACK
HOT AM
FLUE GAS RATE: 26 • W SCfM
1400 IMV
GAS PLOW, Vol. *
SCreajB
»2
C02
H20
°2
S02
S03
NO*
NH-j
Million scTh
123
7«.90 7»-95 75.28
11.70 I*.70 11.83
7.25 7.25 6.98
2.80 2.75 2.86
0.30 Trace Trace
0.30 Trace
0.05 0.05 0.05
1-ju
1,620 T
119.8 1«8.
100
0.771
ELECTROSTATIC
PRECPTTATOR
SULFATf
Figure 9- Kiyoura-T.I.T. Process — Existing Plant
-------
The- temperature at which ammonia is reacted with the flue gas (420°-
500°'') is apparently very important. In this temperature range,
crystals of l-3v and aggregates of >100u of pure ammonium sulfate,
suitable for bagging and shipping, are produced.
Figure 9 shows a flow diagram of a low temperature effluent version
of tlie Kiyoura process. The material balances for both high and
low temperature effluent applications are the same. The major
difference in cost for these versions was due to the requirement
for a flue gas heater for the low temperature application.
The process has essentially the same advantages as the Monsanto-
Penelec process, in that no moving parts are required, the tempera-
ture of the flue gas remains essentially unchanged, there is no
problem of recycling an adsorbent and the sulfur trioxide is easily
recovered as (NH(4)2SOit.
In addition to these advantages the method avoids the expense of a
corrosion problem associated with acid condensation and allows ship-
ment of a dry product rather than a liquid one.
Among the disadvantages of the process are: (1) the requirement for
an additional precipitator (or a bag house), (2) either discharging
the flue gas at (or about) ^00°F or adding another heat exchanger,
and (3) the cost of ammonia and its system. Although it is not men-
tioned, it appears that very precise control of the ammonia/sulfur
dioxide reaction is necessary to prevent, or minimize, escape of
either reactant. The process also creates the problem of product.
marketability in the U. S. where ammonium sulfate has a low market
value and demand.
c. Bayer Double Contact Process (Ref. 103, 104, 105)
The Bayer double contact process is primarily a means of increaring
the overall conversion in a commercial, contact sulfuric acid plant.
Normally, in a contact sulfuric acid plant, conversion .'. s 97-98%,
with the residual 2-3% unconvertedS02 exhausted to the atmosphere.
In the double contact process, conversion is carried to about 90% at
at which point the product S03 is absorbed from the gas mixture.
This effects a shift in the mixed gas composition away from the
near-equilibrium condition, and consequently enhances the conver-
sion rate in a final contact pass. The process not only achieve.*
overall conversion of 99-5$ but, in so doing, virtually eliminate:;
S02 emission from this source. The achievement is costly, however,
because of an additional absorption step and the requirement to
reheat the gas stream for the second contact pass. It is noted
that the economics are favorable only where the price of sulfuric
acid justifies the added cost and that even in such circumstances
the economics become marginal when S02 concentration in the feed
scream drops to about 9%-
57
• MONSANTO RESEARCH CORPORATION •
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It appears that the double contact process would not be applicable
to removal of S02 from flue gas and would be of no benefit even in
the case of a 5% S02 smelter gas. The "breakeven" concentration
of S02 for double contact operation is 6%.
3. Carbon Based Processes
a. Reinluft Process (Ref. 96, 97, 98:, 99, 117, 120, 123)
The Reinluft process for removal of sulfur oxides from power plant
stack gases was invented several years ago at Reinluft Gmbh, Essen,
Germany, by Dr. F. Johswich. The dry process utilizes activated
char to catalyze the oxidation of sulfur dioxide to sulfur trioxide
at a relatively low temperature. The char acts as an adsorbent for
the sulfuric acid that forms from the reaction of the trioxide with
moisture in the gas stream. At high temperature, the sulfuric acid
dissociates into water and sulfur trioxide. The sulfur trioxide is
reduced to the dioxide by reaction with the carbon, thus reactivat-
ing the adsorbent. The desorbed sulfur dioxide may be converted
into elemental sulfur or sulfuric acid.
The Reinluft process has been at a more advanced stage of development
than any other known dry process. The Volkswagen works at Wolfsburg
operated a small plant for several years and were the first to acquire
one of the three prototype units. A second plant was built at the
Carbosulf Company at Cologne. The 1.9 million cu ft per hr plant was'
designed to run on a mixture of- sulfuric-acid-plant tail gas and
waste gas from a Glaus-process kiln. The third unit is the Luenen
installation in November 1966 alongside an experimental power station
operated by Steinkohlen Electricitat AG. It is designed to treat
about 1.15 million cu ft per' hr of flue gas containing up to about
0.9 grain/cu ft of coal dust. The operating level of the Luenen
unit corresponds to that of a power station of about 11-Mw output,
which is about the level of the unit at Cologne also.
In England, active interest has been shown in the process. The
Department of Scientific and Industrial Research has purchased a
pilot-sized plant of the Reinluft design to treat 16,000 cu ft/hr
of flue gas from an oil-fired furnace at their Warm Spring laboratory.
The system consists of a two-stage adsorber section located
over a desorber section, as shown by the flowsheet in Figure 10.
The adsorbent enters at the top of the second stage and moves down-
ward against the gas stream at a rate of 1 to 2 mm/mln. The temper-
ature in this larger stage of the adsorber is about 220°F. Flue
gas, at 250° to 320°F, enters the unit via the first (lower)
stage which acts as a primary high-temperature stripper to remove
all sulfur trioxide and sulfuric acid initially present in the
gas. Since.the oxidation of sulfur dioxide is accomplished at a
relatively low temperature, the gases are drawn off, cooled to
220°F and returned to the adsorber, entering at the second stage.
The clean gas then emerges at the top of the unit, having lost up
to 99.9? of its original sulfur dioxide content. Meanwhile, the 10
58
• MONSANTO RESEARCH CORPORATION •
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FLUE GAS
I
ADSORBER
SECOND STAGE
FIRST STAGE
ADSORBENT
216'P/
® I
SCRUBBED GAS
TO STACK
280-F
CO SOOf
700*1=
BLOWER
REGENERATOR
OESORBMO GAS
VIBRATING SCREEN
FINES
7,300lb/hr
SO, TO
SULRJRIC ACID PLANT
1.2BO ton/dny
(100% ACID)
CHAR MAKE-UP t4,600lb/hr
Scream
Million
N2
C02
H20
Oj
SO,
NO,
SO,
SCPH
Gas
1
71.
11.
7.
2.
0.
0.
150.
Plow,
9
7
25
8
3
05
00
2
71
11
7
2
0
0
150
Volume %
9
7
25
8
3
05
00
3
75.02 .
11.73
7.33
2. 87
Trace
0.05
H9.57
«
26
.18
27
1
25
--
Ij
.05
.10
,80
.05
.00
—
.20
5
.78.05
18.10
27.80
1.05
25.00
iili
6
75
11
7
2
02
73
33
87
Trace
0
2_
05
_6
Figure 10. Reinluft Process for S02 Removal
59
• MONSANTO RESEARCH CORPORATION •
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to 12 mm lumps of adsorbent, loaded with sulfuric acid, move down
through the adsorber Into the desorptlon section. Here the temper-
ature Is raised to 700° to 750°F. Heat is supplied by product gas
circulating at 300°F from near the top of the regenerator, through
a heater, and into the bottom of the desorber (regenerator). In
this section, the sulfuric acid dissociates into S03 and water, and
the SO3 reacts with the carbon to form the product gas, C02 and S02.
A side stream of the product gas containing approximately 25% S02
is withdrawn for subsequent processing. It is not necessary to
place the regenerator under the adsorber. In another arrangement,
the regenerator could be alongside.
In the regeneration, other sulfur-bearing material likewise reacts
to form sulfur dioxide. At this temperature level (700° to 750°P),
side reactions forming carbon monoxide or undesirable sulfur com-
pounds do not occur and the bond between the activated carbon and
sulfur dioxide is so weak that a comparativley small stream of
scavenging gas (C02 or N2) is sufficient to flush away the dioxide.
Reduction of the S03 to S02 during desorption acts to regenerate the
slowly moving charcoal bed. Char is added to the adsorbant leav-
ing the desorber to make up for losses by reaction with S03 and
production of fines. Before recycle, the adsorbant is screened to
remove the fines.
The process uses low-temperature coke as the adsorbant instead of
activated carbon. The coke is formed by vacuum-carbonizing of
coal, lignite, peat, wood, or tarry material at about 1100°F.- Al-
though the raw coke is not "activated," single or repeated im-
pregnation with sulfuric acid and subsequent evaporation at the
high temperature turns it into an effective adsorbant material. 'At
the treatment plant, the low activity fresh coke is used as the
make-up. After three to ten cycles, the activity of the adsorbant
reaches a maximum corresponding to that of the best gas-adsorption
charcoal.
The process has several advantages. Gas cooling is not excessive,
corrosion-resistant equipment is not required, and the by-product,
S02, may be used either to produce sulfuric acid or elemental
sulfur. Reduction of the trioxide consumes part of the carbon
and renders the granules porous. This is equivalent to activation
in place so it mitigates the relatively high cost (cost of low-
temperature coke in Germany is $25 to $90 per ton; activated car-
bon is $750 to $1250 per ton) of activation beforehand. However,
because part of the .carbon is consumed, it must be replaced. Make-
up requires about 0.2 Ib of adsorbant per Ib of sulfur adsorbed.
Among the disadvantages of this process is the cost of recirculat-
ing the very large amount of carbon required, as well as, the cost
of heat and reducing agent. A very prominent disadvantage is the
instability of low-temperature coke in the presence of flue gas
oxygen. Startups were frequently beset by uncontrollable oxidations
and the development of hot spots. This was particularly noted in a
large-scale plant at Cologne which resulted in the redesign of ad-
sorber internals for a unit at Luenen (near Dortmund) to prevent
60
• MONSANTO RESEARCH CORPORATION •
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heat accumulation. As an additional safety measure, the Luenen
plant also had a large vessel installed into which all the activat-
ed coke can be dumped if there is a tendency to overheat. At first,
it appeared that the problem might be overcome, but according to our
latest information the process has been abandoned both in England
and Germany.
b. Sulfacid Process (Ref. 97, 121, 125)
The Sulfacid process was developed by Lurgi (Lurgi Gesellschaft fur
Chemie und Huttenwesen mbH, Frankfurt) in Germany for removal of
sulfur dioxide mainly from flue gases other than power station gases.
The principal difference between this and the Reinluft process is
the method of regenerating the carbon catalyst. The Sulfacid pro-
cess washes the catalyst, to remove the by-product sufuric acid.
The process has been tested in two plants, operating downstream of
oil-fired boilers with flue gas rates of 35,000-53,000 cu ft/hr,
respectively. Two commercial plants are in operation which remove
more than 90% of the S02 from detergent sulfonation waste gases at
flow rates of 35,000 and 88,000 cu ft/hr, respectively. Also oper-
ational are two plants utilizing flue gases containing 4 g S02/cu ft
from oil-fired equipment with capacities of 53>000 to 105,000 cu ft/hr
The process is scheduled for treating the waste gas from a 300-ton/day
sulfuric acid plant, and plans are underway for testing the pro-
cess in a coal-burning power plant as well.
The conversion of sulfur dioxide to sulfur trioxide and sulfuric
acid takes place in a fixed, carbon-containing catalyst bed below
220°F when cooled flue gas is passed through (Figure 11). Sulfur
dioxide is adsorbed on the catalyst where it reacts with oxygen ad-
sorbed from the gas stream and then moisture to form sulfuric acid.
The acid formed in the catalyst pores is continuously rinsed out by
a water spray. The clean, cooled flue gas is exhausted to the atmo-
sphere .
Acid obtained from washing the catalyst bed is about 10-15? H2S04.
It is concentrated during the cooling of the entering flue gas
from about 300°F to the reactor temperature (li»0°-l60°F). In a
packed-tower with counterciirrent" 'scrubbing, the acid concentra-
tion can be increased to 60-70$, depending on the flue gas tempera-
ture. However, when a large volume of gas is to be processed,
as from a power plant (or when the gas is rich in fly ash), a
Venturi scrubber is required for cooling. In this case, the maximum
attainable acid concentration is only 25-30% due to the co-current
operation of the scrubber.
The catalyst mass is essentially activated carbon, although special
additives have been used to accelerate oxidation of the sulfur di-
oxide. Formation of sulfuric acid in the catalyst apparently acts
to clean the bed, as no loss in activity of one bed was noted after
three years of treating the stack gas of a sulfuric acid plant.
The Sulfacid process is simple and avoids the fire-hazard disadvan-
tage of the Reinluft process. However, there is considerable cool-
61
• MONSANTO RESEARCH CORPORATION •
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O
•z.
3)
PJ
m
>.
3)
O
I
n
o
3J
T)
O
a
O
Z
CLEAN GAS TO STACK
140°-160<>F
FLUE GA&
FROM DUST COttECTO*
REACTOR
AT l JL I 1 1
SCRUBBER
DROP CATCHER
JL JL X
H,SQK>-T5%
FEED TANK
S
I ACID PURIFICATION I
j AND CONCENTRATION j
T
Figure 11. Lurgi Sulfacld Process for S02 Removal
-------
ing of the exhaust gas (to 1^0°-l60°P)> and the wet, fixed cata-
lyst beds create a considerable pressure drop in the gas stream.
The sulfuric acid product requires corrosion-resistant construc-
tion materials. Scrubbers and packed towers for cooling and con-
centrating, and the converter, are reportedly of mild steel with
lead or rubber lining.
c. Hitachi Process (Ref. 97, 126)
A process similar to the Sulfacid process is under development by
Hitachi Ltd. at the Goi plant of Tokyo Electric Power Company
where a 2 Mw test unit is in operation.
The system (Figure 12) consists of six towers charged with
activated carbon catalyst. The towers operate in a cyclic man-
ner: a single tower goes through a cycle of 30 hrs of adsorption
of SO2 from uncooled (300°P) flue gas from the dust collector, 10.;/«
hrs of washing (during which, no gas passes through the tower), ;'
and 20 hrs of drying. At one time, three towers are adsorbing S02,
two towers are drying and one tower is being washed. Therefore,
the gas flow in the system is changed every ten hours, and the
total time of one cycle becomes 60 hrs.
After 30 hrs of dry adsorption, the gas flow through a tower is
stopped and water enters the top of the tower to wash the sulfuric
acid from the carbon. The acid, formed directly on the carbon
during the adsorption period is about 70% l^SOj,. Acid concentra-
tion in the wash, however, is only 10-20% H^SOitj and even this
level is attained only through a type of staged, counter-current
washing operation.
Wet carbon, previously washed to remove sulfuric acid, contains .
20-50 wt % water and is dried by hot flue gas. The temperature
of the gas in the drying section drops to 120°-l60°F until the car-
bon is dry, at which time the gas temperature rises to that of the
entering flue gas, ca 300°F. There is some adsorption of S02 during
drying so that the net total adsorption time in a 60-hour cycle is
50 hours.
Qas from the drying section is mixed with flue gas from the feed blower
and enters a dry adsorption tower where S02 and moisture from the
drying section are removed. The humidity of the gas stream to the
atmosphere is thus very near that of the flue ga? from the boiler.
Heat of adsorption and reaction result in a temperature rise of
10°-30°F during adsorption to give a net temperature drop in stack
gas temperature of only about 55°F.
63
• MONSANTO RESEARCH CORPORATION •
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GAS TO STACK
2
O
z
>
z
O
a
01
(A
cr>
a
o
i
o
o
X
~Q
O
O
Z
FLUE GAS
FROM DUST COLLECTOR
A JOO°F
H,0
1,0
CONCENTRATION
WASHING TANKS
Figure 12. Hitachi Process for S02 Removal
-------
The equipment for a process plant would include the following:
Six towers containing activated carbon.
Feed blower.
Circulation blower.
Pump for wash water.
Purge pump.
Tanks for storing washings.
Acid concentration equipment.
Dampers for changing gas flow from tower to tower.
The process has the advantages and disadvantages of the Sulfacid
process. Investment costs in corrosion-resistant materials and
dampers would be high; the acid concentration is low and would
involve increased cost for processing. Some pressure drop in
the fixed bed would occur; however, compared with Reinluft, the
process is safe, and the effect of the process on stack gas temper-
ature is small; the exhaust temperature of the gas is over 220°F.
The process is simple in principle, removal of the H2S0lt is easy,
and there appear to be no drying problems. Furthermore, the pro-
cess removes more than 90? of the S02 from power plant stack gas.
^. Manganese Based Processes
a. Mitsubishi Process (Ref. 97, 110, ll^J)
Mitsubishi Heavy Industries Ltd., Tokyo, has a process for removal
of sulfur dioxide from flue gases that relies on the reaction of
sulfur dioxide with manganese dioxide to form manganese sulfate.
Regeneration of the manganese dioxide is accomplished by reaction
of the sulfate with ammonia. Ammonium sulfate is produced as a
by-product.
Mitsubishi originally tested the process at the Yokkaichi oil-
burning power station of Chubu Electric Power Company. The small
test unit treated gas equivalent to about 1 Mw. Following the
success of the pilot plant study, a large desulfurization unit was
built at the same location to process 25% of the flue gas output of
the 220 Mw oil-burning power plant.
The Mitsubishi process (Figure 13) is composed of the following
three major units:
A. Gas adsorption units consisting of adsorption tower,
multiclone, electrostatic precipitator, blower and
absorbent feeder.
B. Regeneration units consisting of oxidizing tower,
ammonia scrubber, air compressor and filter.
65
• MONSANTO RESEARCH CORPORATION •
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VACUUM
v>
z
o
a
n
m
o
x
o
o
a
TJ
o
FLUE CAS FROM
AIR PREMEATEK
300»F
Figure 13. Plow Diagraii of Mitsubishi Manganese Dioxide
Process
-------
C. Crystallizing units consisting of reactor, crystal-
lizer and dryer.
Flue gas from the air preheater enters an adsorption tower into
which active powdered manganese dioxide is uniformly dispersedt
The mov.lng bed adsorber is said to be suitable for large volumes
of flue gas containing low amounts of sulfur dioxide. The linear
velocity of the gas is increased by-about 50% in the adsorber.
The temperature of the gas remains between 212° and 356°F. The
amount of adsorbent used in the unit is 150 to 250 g/cu m.flue gas.
In the adsorber, sulfur dioxide reacts with manganese dioxide to
form manganous sulfate. The exit gas enters a cyclone, where
about 90% of the unreacted manganese dioxide, and the manganese
sulfate, are collected and returned to the gas stream entering
the adsorption tower. The residual 10% of the solids mixture is
collected in an electrostatic precipitator. The adsorbent exists
as relatively large 40y high density particles. From 203° to 257°F
the resistivity of the adsorbent is between 3 x 10~7 and 8 x 10"7
ohm-cm. This and the size allow trouble-free collection in the
electrostatic precipitator.
From the precipitator, solids are sent to a tank and slurried with
about 70? water. The slurry is passed through an ammonia recovery
tower and into the regenerator (oxidizing tower). Air and ammonia
are injected into the regenerator where hydrated manganese oxides
and ammonium sulfate are produced. The solid oxides are separated
from the ammonium sulfate solution by filtration and returned to
the flue-gas inlet. The ammonium sulfate solution goes to a crystal-
lizer for recovery of the solid product.
For an oil-burning power plant, flotation with kerosene was used
to remove soot or ash from the slurry at the soot separator. This
was found to lengthen the useful life of the adsorbent and improve
the purity of the ammonium sulfate.
There are several advantages to the Mitsubishi process. Because it
operates at stack gas temperature, it is easily adapted to existing
power plants. The exhaust gas is still dry and' warm after treatment
providing favorable plume characteristics. Mild steel fabrication
is indicated throughout, except in slurry treating units. Low pres-
sure drop in the adsorber suggests economical movement of gas through
the treatment unit.
Among the disadvantages of the process is the fact that the by-
product, ammonium sulfate, has essentially no market potential in
this country. In Japan, or other parts of Asia, the product market
potential may be brighter. Though simpler than most systems for
this type of process, the manganese dioxide regeneration is still
67
• MONSANTO RESEARCH CORPORATION •
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complex. As the station size increases, the materials handling pro-
blem associated with the adsorbent also increases until, at the
1*100 Mw level this amounts to moving 16 tons/min of adsorbant.
b. TVA Ozone-Manganese Processes (Ref. 106, 107, 109, 112)
In the process variations of direct acid process and direct
ammonium sulfate process, manganous ion appears to be acting as a
true catalyst, However, unless ozone acts as a separate catalyst,
its obscure role in conjunction with manganous ion might be con-
sidered unnecessary.
In the direct acid process, (Figure 14), ozone, at concentrations
up to 240 ppm, is injected into the flue gas stream. The gas is
then cooled to 130°-170°F and scrubbed in a packed tower with a
liquor consisting of sulfuric acid and manganous ion, the latter
in concentration from 0,02? to 0.2%. In the scrubbing tower, sul-
furic acid is produced at a maximum concentration of 40%. The
weak acid can be concentrated in cooling the entering flue gas to
scrubbing temperature. No provision is described for removing
manganous ion from product acid. Conceivably, this may not be
important to product use, but it represents considerable loss of
manganese.
In the direct ammonium sulfate process, Figure 15> the scrubbing
liquor consists of ammonium sulfate solution and manganous ion in
the same concentration as above. Operation is essentially the
same, except that the sulfuric acid formed is continuously
neutralized with ammonia. The ammonium sulfate solution can be
concentrated, to some extent, In cooling the entering flue gas;
but, to produce a dry product, considerable additional equipment
would be required. Again, there is no description of manganous
ion removal from the product.
In both processes, retention times in the scrubbers ranged from
18 to 36 seconds, with better performance at higher retention
time. The major part of TVA's study was done with simulated flue
gas consisting of 0.35/5 S02, 3-5% 02, 16% C02 and 80% N2. The
composition corresponded approximately to gas produced by burning
coal containing 3.5% sulfur with 10% excess air. In tests with
flue gases from a power-plant and a pilot plant coal burner, the
process was far less efficient; recovery of a high percentage of
S02 required a retention time on the order of 90 sec. The ef-
ficiency of ozone utilization, even with a retention time of 90
sec, was very low.
Apparently, some component in the combustion gas poisons the cata-
lysis. Thorough removal of particulate matter did not improve
68
• MONSANTO RESEARCH CORPORATION •
-------
formance indicating the poison(s) to be gaseous. The identity of
the poison(s) was not determined; however, phenol, sometimes found
in coal combustion gases, is known to poison manganese catalyst.
These processes are saddled with far more disadvantages than ad-
vantages, perhaps the greatest disadvantage in relation to large
power stations, is the long residence time of the gas in the
scrubber.
5- Selenium-Based Processes
Recently, two processes have been described for removal of S02
from flue gas based on the stoichiometric reaction
2S02 + Se02 + 2H20 > Se + 2H2SOit (28)
This wet-process reaction proceeds at a rate roughly 10 times, the
rates obtained with wet reactions employing elemental oxygen cata-
lyzed by ozone, manganese ion, or carbon. The possibility of com-
pletely and economically removing S02 derives from the fact that
the selenium is precipitated in solid form and is insoluble in
the product acid. The reaction is therefore irreversible and pro-
ceeds to completion. Elemental selenium is readily filtered out
permitting recovery of a "clean" product acid, a distinct advantage
over the manganese wet processes. Selenium is converted back to
the dioxide by "burning" it in air. Since the dioxide sublimes at
about 350°C, a means is thus afforded of separating the material
from residual fly ash contamination.
a. Nor Deutsche Affinerie Process (Ref. 127)
In the Nor Deutsche process, Figure 16, flue gas passes through
two scrubbing towers sequentially. In the second tower, the
gas is scrubbed with a solution of selenous and sulfuric acids
in which the concentration of selenium dioxide is maintained con-
stant and in excess of that required to oxidize the S02, some of
which was removed in the first tower. The exhaust gas from the
second tower is essentially free of S02 and vents to the stack.
Selenium is continuously filtered out of the recycle scrubbing
acid, oxidized, and redissolved in the scrubbing acid. Aliquots
of the scrubbing solution are withdrawn at intervals from, the
second tower and sent to the first tower where recycle continues
until all of the selenium dioxide is reduced by incoming flue gas
to selenium-free product. Product acid leaves the system through
^concentrator heated by incoming flue gas. It is claimed that 60°
Be sulfuric acid is attainable in this manner.
69
• MONSANTO RESEARCH CORPORATION •
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TO STACK
M8.2xlO*»cfh
VAPOI TO
OZONE
3000* cfh
FLUE GAS FROM
AIR PREHEATER
150xl06»cfh
;-j
1°
o
i
3
X
5
MANGANESE SULFATE
735lb/hr
Figure 14. Flow Diagram of TVA Direct Sulfurlc Acid Process
-------
TO STACK
MIST
fUMINATOI
|
o
en
m
O
X
o
o
•a
o
FLUE GAS FROM
AIR PREHEATER
WATER
TO ECONOMIZER
AMMONIUM
SULFATE
Figure 15. TVA Direct Ammonium Sulfate Process
-------
w
I
H
•a
C!
m
>
o
i
3
X
6
"*
FLUE GAS
0.3% S02
H,S04
1
CONCENTRATOA
1
BYPRODUCT
H,S04
I75°F
••••
1 F
STORAGE
TANK
1 1
1
i
\
i
h
A
7
s.
H,S04
FILTER
1
MM
i
t
SCRUMING
TOWERS
1 i
STORAGE
TANK
1 1
1
\
<
i
A
s.
S.O,
HzSO
i
FILTER
masi
$.0, H,S04 J
S. "
<
>•
•
i
C
1
S«0j
SELENIUM
•URNER
CLEAN GAS
TO STACK
t
AIR
(OUTSIDE THE SYSTEM)
Figure 16. Nor Deutsche Affinerie Process for S02 Removal
-------
b. Badische Anilin- and Soda-Fabrik Process (Ref. 128)
The Badische process, Figure 17, employs a single scrubbing tower,
or a venturi scrubber, in which SC>2 is oxidized by selenium dioxide
in a recirculating solution of selenous and sulfuri'c acids. The
selenous acid concentration is maintained constant and in excess
of stolchiometry with solid selenium continuously filtered out, re-
oxidized and reintroduced into the system from .a side stream, In
this process, residual selenium dioxide'.is reduced in the main pro-
duct stream from the scrubber by reaction with another source of S02,
e.g. roaster gas. The excess S02 in the gases from this, step is vent-
ed into the Incoming flue gas upstream of the removal system along
with reoxidized selenium in vapor form.
The oxidation rate of SC>2 in the scrubbing liquor decreases with in-
creasing concentration of ^SO^. However, the process may be operat-
ed in a multistage-mode with each stage containing successively high-
er concentrations of H2SOi, in all of which Se02 is quite soluble.
For example, in a two-stage process, the first stage liquor may be
selenous acid and 30% H2SO^. The second stage would scrub the flue
gas wi.th H2SeC>3 in 30% H2SO(t. The H2SO», produced in'the first stage
would be replaced by the 30% H^O^ from the second stage; A three-
stage system would run with first stage acid concentration at 70%,
while the third stage would use 10% acid with an excess of selenous
acid.
The basic difference in the two selenium processes lies in the mode-
of operation to produce a ."clean", marketable product. The Nor
Duetsche process strips selenous acid from the scrubbing liquor with
excess S02 in the flue gas and concentrates the product acid with
heat in the flue gas. The Badische process uses an outside source of
S02 to strip selenous acid from the scrubbing liquor, witfh product
concentration being dependent upon the number of scrubbing stages.
The selenium processes appear to be relatively simple. The oxidant
is recoverable, leading to low consumption'and therefore minimum
oxidant make-up. The production of marketable sulfuric acid is an
advantage. Reduction in equipment size, compared with other wet
processes, is indicated by the relatively high rate of oxidation of
S02 by Se02.
The processes, however, also have some obvious disadvantages.
Selenium is expensive and its oxides are highly toxic and corrosive.
Removal of the elemental selenium from the system for reoxidatlon
to Se02 and re-introducing it to the system could present a handling
problem. Also, additional equipment and facilities would probably
be required to wash and dry the selenium residue to remove absorbed
sulfuric acid and separate the new Se02 from accumulated fly ash
that was not retained by the dust collection system. Because the
73
• MONSANTO RESEARCH CORPORATION •
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FLUE GAS 300°F 0.3% SO2
VENTURI
SCRUBBER
CLEAN CAS
TO STACK
i
MIST
ELIMINATOR
(OUTSIDE THE SYSTEM)
StOj
AIR
H2S04
BYPRODUCT
Figure 17. Badische Anllin- and Soda-Fabrik Process
for S02 Removal
• MONSANTO RESEARCH CORPORATION •
-------
liquor and by-product is sulfuric acid, corrosion-resistant equip-
ment would be required throughout the system. The temperature of
the "cleaned" flue gas when exhausted to the atmospher is lower
than needed to produce a high plume. However, since the reaction
of Se with 02 is exothermic, it may be possible to raise this temper-
ature somewhat with the heat generated, by reoxidizing the elemental
selenium.
Apparently, bench scale work has been done to evaluate some operat-
ing parameters. However, there is no knowledge that any pilot plant
studies have been conducted, or are anticipated. This quite possibly
may be attributed to the toxicity and cost of selenium.
6. Modified Chamber Process
Tyco Laboratories, Waltham, Massachusetts, is studying a modified
chamber-process approach to flue gas cleaning with NAPCA support.
Since the oxides of nitrogen in flue gas are also objectionable, it
is desirable to remove them along with the sulfur. Further, the oxides
of nitrogen are capable of catalyzing the oxidation of sulfur dioxide.
The obsolete chamber process utilized this principle to produce sul-
furic acid commercially for a number of years.
In an early version of the Tyco process, Figure 18, flue gas at about
300°F enters the reactor where SO^ is oxidized by nitrogen dioxide at
a mole ratio of 3N02:NO:S02. Retention time in the reactor is esti-
mated at one to two seconds. Sulfur trioxide, nitrogen oxides, and
water vapor are absorbed from the cooled reactor effluent in 80$
sulfuric acid to form nitrosylsulfuric acid. Nitrogen oxides are
stripped from the nitrosylsulfuric acid, sent to an oxidizing chamber
where the proper mole ratio of N02:NO is re-established, and then
recycled to the converter. Since the entering flue gas contains
some nitrogen oxides, there is a net accumulation of these in the
system, and a side stream is used to make nitric acid, a second by-
product of the process. Eighty percent sulfuric acid is obtained
as by-product in a side stream from the absorber liquor.
Since the process is still in a very early stage of development, it
is difficult to assess accurately the advantages and disadvantages.
However, as initially described, the process clearly has one major
advantage and one major disadvantage: it removes S02 and some NOX
at low temperature, but it requires the power station to burn ^0%
more fuel to supply high process heat requirements. Tyco is build-
ing a 10 c.f.m. pilot plant to investigate methods of lowering
this prohibitively high heat load. Recent results indicate a
more favorable heat balance.
Heat and material balances for an early version of the Tyco process
are summarized in Figure 18. A later version of the process is
discussed further in Volume II. In addition to general assumptions,
the calculations here were based on the following assumptions:
75
• MONSANTO RESEARCH CORPORATION •
-------
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Ab«arb«r
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1 33O"F
(T>
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Flue Gas Fate: 150 x 106 SCFH
Cas Flow, Welp.ht I
1567
71.00 70.18 61.75 61.75
22.82 21.61 19.01 10.01
0.36 1.10 10.29 10.29
2.82 3.07 2.51 1-65
trace 0.65
0.06
trace — 3-88 6.36
0.18 trace — 2.53 0.0]
H2SOk
Million
11.810 2.805
8
(to
(JL)
1
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f
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M,80
-------
• S02 removal: 95%
• NOX removal: 95%
• Molar ratio of N02 : NO: SC>2 entering the reactor: 3:1:1
• Additional coal required for stripping gas: 21%
• Nitric acid concentration: 50%
7. Reversible Dry Absorbent Process
Gallery Chemical Company, Division of Mine Safety Appliances
Company, has considered a system for removing and concentrating
the product SOs after conversion of the SOz in flue gas. A pre-
liminary process design resulting from this NAPCA supported study
is shown in Figure 19.
The exhaust from a catalytic converter is passed through a bed
of absorbent (20 wt % of Na2SOi+ on inert support) where product
SOa is absorbed (forming Na2S20?). Exhausted absorbent is re-
generated with hot gas which releases the S02. The regenerant
stream containing 20 vol % S03 Is fed to an acid absorption column
to produce 100% acid or oleum.
i
For a fixed bed system, four absorption columns are required to
permit adequate cycle timing, as illustrated by the following:
ABSORBER NUMBER
15 min. abs.
15 mln. des. 15 min. abs.
15 mln. des. 15 min. des. 15 min. abs.
15 min. cool. 15 min. des. 15 min. des. 15 mln. des.
This process is described in greater detail in Volume II.
Advantages of the process include the production of 100% acid, and
elimination of the need for corrosion resistant equipment in the flue
gas stream.
Disadvantages include effects of thermal cycling on absorbent, unknown
absorbent attrition rate, and early development status;
77
• MONSANTO RESEARCH CORPORATION •
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Plue Ca> Rate: 150 > 10' SCPH
Gai Flo., Ml. f
o
z
2
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O
m
m
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o
o
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X
H
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•
co
CO,
H,O
o,
so,
so,
•0,
B,30»
(A)
Flue Gas From
70.18
21.6*
*.«0
2.93
0.06
0.73
0.06
70.69
21.79
2.96
0.06
Trace
0.05
c
70.18
21.6*
«.«0
3.07
0.65
0.06
• 1.60
12.82
2.61
1.82
0.38
»0.7»
0.03
E
70.19
21.6«
«.»0
3.08
0.61
Trace
0.05
0.5
70.69
21.79
2.95
0.07
Trace
0.05
11.810 11.729 0.118 0.199 0.118 0.100 11.8*7
Flue Gas To
Converter
JJOLE
Cooverter 8MfF
Absorber-Stripper
lOHfl
Gas
Cooler
Air
©p—|
Furnace
Converter
Pre-heater
Flue Gas" From
Precipitator
450-500°F u
cn
503
&KcArhor
^
^-*
L-«^
Mist
Eliminate
_j
ACID
PROOUQ
Acid
Cooler
Figure 19. Gallery Dry Absorbent Process
-------
8. Process Cost Estimates
Table 10 summarizes the capital requirements and operating costs
for the 7 processes. Capital requirements for Sulfacid and TVA
direct acid process were based on limited literature data. Esti-
mates for Hitachi, TVA-ammonium sulfate, selenium and Tyco processes
are not included because of insufficient design information on
these processes. Preliminary capital requirement estimates and
operating costs were based on the following assumptions:
General:
• Flue gas rate - 2.5 MM SCFM
• Size of power plant - 1*100 MW
• Flue gas analysis:
Component
N2
C02
H20
02
S02
NOX
Fly Ash
by Volume
7^.9
14.7
7.25
2.8
0.3
0.05
0.2 (by weight)
Coal required - 580 tons/hr
Operating factor - 330 days/yr @ 100% capacity
Direct labor - $3.00/hr
Supervision - $7,800 and $12,000 annually for first-
line supervisors and area superintendents, respectively
Maintenance - 5% of the fixed capital investment
Plant supplies - 15% of the maintenance cost
Utilities:
a) steam - 50
-------
Table 10
Capital Requirements and Operating Costs
for S02 Oxidation Processes
(for capital cost breakdown, see Appendix IV)
2
O
z
(/>
z
m
m oo
> o
ro
o
i
o
o
71
TJ
O
X
>
H
O
Z
Operating Cost
Process
Monsanto-Peneleo 1
Monsanto-Penelec2
Kiyora-T.I.T.1
Kiyora-T.I.T.2
Reinluft1
Sulfacid1
Mitsubishi1
T.V. A. -Direct Acid1
Gallery Dry Absorb-
ant*
Capital Requirement
Dollars $/KW
37,189,
3^,067,
25,889,
22,473,
31,160,
67,925,
22,869,
61,600,
000
000
000
000
000
000
000
000
16,830,000
26
24
18
16
22
48
16
44
12
.77
.33
.49
.05
.47
.51
.33
.00
.02
$/Ton
$/Yr. of Coal Mills/KWH
10,643,000
9,506,000
16,568,000
15,425,000
13,531,000
24,320,000
16,450,000
21,536,000
4,724,150
2.
2.
3.
3.
2.
5.
3-
4.
1.
32
07
60
36
95
29
58
69
03
0.
0.
1.
1.
1.
2.
1.
1.
0.
96
86
49
39
22
19
48
94
43
By Product
70 to 80? sulfurlc
70 to 80% sulfuric
Ammonium sulfate
Ammonium sulfate
985? sulfuric acid
93% sulfuric acid
Ammonium sulfate
80X sulfuric acid
9955 sulfuric acid
acid
acid
1. low temperature effluent (about 300°P) from existing power plant
2. high temperature effluent (about 850°F) from new power plant
3. low temperature effluent from existing power plant, including catalytic converter
-------
c) water
raw (untreated) 104/1000 gal
process (treated) 254/1000 gal
recirculated cooling 54/1000 gal
tower
d) compressed air - 54/1000CF
• Payroll burden - 2055 of direct labor and supervision
• Plant overhead - 50% of labor, maintenance and supplies
• Depreciation - 10% of the fixed capital investment
• Taxes - 2% of the fixed capital investment
• Insurance - 1% of the fixed capital investment
• Working capital - 10$ of the fixed capital investment.
Monsanto-Penelec Process
• S02 conversion: 90%
• S03 recovered as H2SOJt: 95%
• Contact time: 0.3 sec.
• Pressure drop across catalyst bed: 4 in H20/ft of depth
• Vanadium catalyst life: 5 years
• Vanadium catalyst cost: $lo45/liter
• Density of vanadium catalyst: 36 Ib/ft3
Kiyoura - T.I.T. Process
Same as for the Monsanto-Penelec process, plus the following:
• Ammonia cost: $60oOO/ton
• Size of (NHit)2 SOi, particles: 1-3 microns,
Aggregate size: above,100 microns
Reinluft Process
• Adsorption of S02 by carbon: 0.1 Ib S02/lb carbon
• Carbon losses:
Chemical.Reaction: 0.2 Ib/lb of sulfur
Attrition: 0.2 Ib/lb of sulfur
• S02 recovered: 9555
• S02 concentration in product gas: 25%
• Conversion efficiency for sulfuric acid plant: 96%
81
• MONSANTO RESEARCH CORPORATION •
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• Residence time in adsorber: 1 sec.
• Linear velocity in adsorber: 100 ft/sec
• Adsorbent carbon cost: $80.00/ton
• Vanadium catalyst cost for sulfuric acid plant: $l.'J5/liter
• Vanadium catalyst life for sulfuric acid plant: 5 years
Sulfacid Process
• S02 conversion: 90$
• Contact time: 10 sec
• Linear velocity in reactor: 100 ft/sec
• Catalyst life: 3 years
• Cost of active coal: $220.00/ton
• Density of active coal: 22 lb/ft3
Mitsubishi Process
• S02 conversion: 9Q%
• Absorbent used: 200g/cu m
• Cyclone efficiency: 90$
• Amount of water in slurry: 70%
• Residence time in absorber: 1 sec
• Linear velocity in absorber: 150 ft/sec
• Absorbent loss due to attrition: 0.1 Ib/lb of sulfur
• Ammonia cost: $60.00/ton
• Absorbent cost: $l44.00/ton
TVA-direct sulfuric acid
• S02 conversion: 90%
• Amount of MiSCv 0.3 Ib MnSOit/100 Ibs water
• Amount of ozone injected in flue gas: 20 P.P.M.
• Residence time in packed scrubber: 10 sec
• Linear velocity in packed scrubber: 70 ft/sec
• MnS04 cost: $95.00/ton
82
• MONSANTO RESEARCH CORPORATION •
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Gallery Process
S02 converted to SOa - 90%
Amount of NazSOi* In total absorbant - 20% (by weight)
100% absorption efficiency, i.e., stoichiometric ab-
sorption, for 15 minutes
Vanadium catalyst cost - $l.'l5/liter
Vanadium catalyst life - 5 years
Cost of Na2SO<4 - $30.00/ton
Cost of Coal - $4.00/ton
Tables 11 through 19 summarize the preliminary cost estimates of
major equipment for the seven processes. As it is possible to have
low and high temperature effluent feed to Monsanto-Penelec and
Kiyoura-T.I.T. processes, cost estimates at both conditions were
made for these two processes. The other processes were based on
low temperature effluent (300°P). Fixed capital costs were esti-
mated after application of Lang's factors of 4.7*1 for fluid pro-
cess plants and 3«63 for solid/fluid process plants. Working
capital was set at 10% of the fixed capital investment for estimat-
ing the total investment.
The operating costs of Monsanto-Penelec, Kiyoura-T.I.T., Reinluft,
Callery, Sulfacid, Mitsubishi and T.V.A.-Direct Acid processes
were estimated. Tables 11 through 19 summarize these estimated
operating costs. The credit for by-product sulfur, sulfuric acid
or ammonium sulfate has a major effect on the net operating costs.
Since the prices of these by-products fluctuate, Figures 19 to
27 present the relationship of sales value of by-product to the
net operating cost. Capital cof-ts ore summarized in Appendix IV.
9. Comparative Evaluation of Flue Gas Treatment Processes
Although a number of processes have been described, only a few have
been well-developed, and, even among these, not all are worthy of
further consideration. The preliminary cost estimates have a
good deal to say in this respect. .It remains to be pointed out
those approaches that might well be pursued and those that should
not, in view of current information.
• MONSANTO RESEARCH CORPORATION •
-------
The wet manganese-based processes do not appear to hold promise. The
wet methods studied by TVA display no characteristics to suggest
further study. The dry manganese oxide, non-catalytic process of
Mitsubishi produces ammonium sulfate, which, it seems, can be pro-
duced as efficiently by the Kiyoura catalytic process. Further,
the Mitsubishi process can produce only ammonium sulfate (or an
alkali sulfate), whereas, the Kiyoura process can produce either
this product or sulfuric acid. The one salient advantage of the
Mitsubishi process is its adaptability to low temperature stack
gas, i.e. existing power stations.
In this latter respect, the carbon processes are tempting to
further consideration. However, associated with the dry carbon
techniques are two grave obstacles: 1) the tenacious bond between
carbon and product sulfuric acid, and 2) the instability of carbon
in the presence of the relatively large amount of residual oxygen
in the flue gas. Both disadvantages are manifest in the Reinluft
process, which, finally, is an adsorption technique producing SO
in a concentrated stream. In attempting to circumvent these
obstacles, the wet carbon processes have introduced a new set of
problems which limits their potential to small and specialized
installations where long retention times in adsorbers and weak
acid product may not count so heavily. Otherwise, there appears
to be no basis for further consideration of carbon.
It has already been Indicated that the selenium dioxide reaction
with sulfur dioxide is not catalytic but stoichiometric. The
health hazard, associated with selenium compounds, has also been
noted. Nevertheless, selenium dioxide presents some interesting
aspects for .potential treatment processes. For example (a) Does
the reaction proceed at a satisfactory rate in the vapor phase?
and (b) Since there is excess oxygen in the flue gas, in a bed of
solid selenium dioxide, could the following reactions be made to
occur concurrently through the bed?
2 S02 + Se02 " 2 S03 + Se (29)
Se + 02 " Se02 (30)
From an academic viewpoint, if the reaction rates are good enough,
the concept under (b) offers interesting process possibilities for
low temperature application.
• MONSANTO RESEARCH CORPORATION •
-------
The Tyco modified chamber-process approach to flue gas S02 oxidation
presents a number of problems. Again, this process offers ready adapt-
ability to low temperature operation - its major advantage. However,
this process is under separate investigation by NAPCA.
Finally, there are the vanadium-based processes. In producing am-
monium sulfate, Kiyoura obviates the high costs of corrosion resis-
tant materials of construction. On the other hand, in this country,
the product would present a considerable disposal problem.
Except for some proprietary engineering and operational differences
we might say that there is only one vanadium process, since either
of the two described could produce sulfuric acid or ammonium sulfate.
The greatest advantage of the vanadia process is its simplicity; the
major disadvantage is its poorer operating economics when applied to
low temperature effluent (existing plants). Further, the Monsanto-
Penelec version of the vanadia process affords, with minor modifica-
tions, the same opportunity for NOX removal as the Tyco (modified
chamber) process.
The Gallery reversible dry absorbent process is not per se an S02
removal process, as it depends upon a prior step to convert S02
to SO,. However, it is presented here because of the apparent ef-
fect it has on improving the economics of a process like the Monsanto-
Penelec process (see Table 21). it also points up the potential
value of a process that could sorb S02 directly from flue gas and
desorb a concentrated stream of S02 to a contact converter for acid
production or to a Glaus reactor for elemental sulfur production.
85
• MONSANTO RESEARCH CORPORATION •
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Table 11
Or.Ki-.ATJNC. COHT ESTIMATE SUMMARY
Basis: 330 Day/Year 6100? Capacity
Category: Existing Power Plant (Low Temperature
Effluent)
Name of ProcessMonsanto-Penelec Flue Gas Rate
2.5
MMSCFM
MW 1400
Fixed Capital Cost $3"
jOBl.OOO
ITEM TOTAL $
1.
2.
3.
4.
5.
6 .
7.
8.
9.
10.
11.
12.
13.
1*4.
15.
16.
17.
18.
19.
20.
21.
22.
23-
24.
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5? of Fixed Capital 1
Supplies, 15? of Maintenance
Utilities 2
Other
TOTAL DIRECT COST _5
Payroll Burden, 20? of 2 & 3
Plant Overhead, 50? of 2, 3,
4 and 5 1
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST 1
Depreciation, IQ ? Fixed
Capital/Yr. 3
Taxes, 2? of Fixed Capital
Insurance, 1? of Fixed Capital
Other
TOTAL FIXED COST 4
TOTAL OPERATING COST \£
COST: $/Ton of Coal 2 32
Mill/kwh 0.96
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh
250,000
105,000
31,200
,700,000
255',000
,800,000
.111,200
27.200
,045.600
,072,800
,408,100
680,000
341,000
.429.100
,643,100
PER CENT
2.35
0.98
0.30
15.98
2.39
26.30
48.30
0.26
9.75
10.01
32.10
6.38
3.21
41.69
100.00
86
• MONSANTO RESEARCH CORPORATION •
-------
in
c--
Q
U
<
U
oe
U
3
Q
O
<*
O
2
O
y
w
a.
20-
15-
10-
ou
1.0
BREAK EVEN POINT
PROFIT
COST
-------
Table 12
OPr.SATlHO COST ESTIMATE SUMMARY
Kasis: 330 Day/Year gl°°iE Capacity
Category: Mew Power Plant (High Temperature
' Effluent)
Name of Process 'lonsanto-Penelec Flue Gas Rate
MW HOP
Fixed Capital Cost$30,970,000
2.5
MMSCPM
1.
2.
3.
H.
6.
7.
8.
9.
10.
11.
12.
13-
11.
16.
17.
18.
19.
20.
21.
22.
23.
214.
ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5* of Fixed Capital
Supplies, !•>* of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20* of 2 & 3
riant Overhead, 50* of 2, 3,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, :. c /• f-'/.xed
Taxes, 2% of Flx-.-7rt •?.-,;..•! 1. •-..•.
Insurance, 1* of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 3.07
Mlll/kwh Q.86
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mlll/kwh
TOTAL $
2:50, COO
105,000
31,200
1,550,000
232,000
2,325,000
1,193,200.
27,200
959,100
986,300
, r..,T cr:,.
•'• :•?,»! oo
309,700
1,026.100
9.505.600
PER CENT
2.63
1.11
0.33
16.30
2 J| 4
21.15
17.26
5729
10.09
10.'38
32.58
3.26
• 12.. 36
100.00
• MONSANTO RESEARCH CORPORATION •
-------
20-
in
c--
o
0
y 15
et
3
U
3
O
O
at
a.
a.
o 10
z
o
at
a.
0.4
1.0
BREAK EVEN POINT
PROFIT
I
MILLS PER KILOWATT HOUR
0.2
0.2
0.4
0.6
0.8
0 1.0
$ PER TON OF COAL
2.0
Figure 20. Effect of Product Credit on Operating Cost
of Monsanto-Penelec Process (new Plant)
89
• MONSANTO RESEARCH CORPORATION •
-------
Table 13
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year 6100$ Capacity
Category: Existing Power Plant (Low Temperature
Effluent)
Name of Process Kiy°"ra - T.I.T. Flue Gas Rate
2.5
MMSCFM
MW 1100
Fixed Capital Cost
$23,535,000
1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19-
20.
21.
22.
23.
21.
ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15? of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20$ of 2 & 3
Plant Overhead, 50$ of 2, 3,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 $ Fixed
CapTtaT/Yr.
Taxes, 2$ of Fixed Capital
Insurance, 1.$ of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 3.60
Mlll/kwh 1.19
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mlll/kwh
TOTAL $
8,150.000
105.000
31.200
1.175.000
176.000
2.800.000
27r200
71s. 600
770.800
2.353.500
170.700
235.350
3.059.550
16,567,550
PER CENT
51.00
0.61
0.19
7.09
1.06
16.90
0.17
1.18
1.65
11.21
2.81
1.12
. 18.17
100.00
90
• MONSANTO RESEARCH CORPORATION •
-------
401
30-
z
o
u. 20-
O
Z
o
u
c*
a.
10
2.0
BREAK-EVEN POINT
PROFIT
0.6 0.4 0.2
COST
MILLS PER KILOWATT HOUR
0.2 0.4 0.6 0.8 1.0 1.2
t i i • i •
1.0
1.0 2.0
$ PER TON OF COAL
3.0
1.6 1.8
i
4.0
Figure 21. Effect of Product Credit on Operating Cost
of Kiyoura-T.I.T. Process (Existing Plant)
91
• MONSANTO RESEARCH CORPORATION •
-------
Table 14
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year 6100% Capacity
Category: New Power Plant (High Temperature
Effluent)
Name of Process Kiyoura - T.I.T. Flue Gas Rate
2.5
MMSCPM
MW 1100
Fixed Capital Cost*20
,130,000
ITEM TOTAL $
1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13-
It.
15.
16.
17.
18.
19.
20.
21.
22.
23.
21.
Raw Materials & Chemicals 8
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital 1
Supplies, 1535 of Maintenance
Utilities 2
Other
TOTAL DIRECT COST 12
Payroll Burden, 20* of 2 4 3
Plant Overhead, 50$ of 2, 3,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed
CapHaT/Yr . 2
Taxes, 2% of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FIXED COST 2
TOTAL OPERATING COST !5
COST: $/Ton of Coal 3-36
Mill/kwh 1.39
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh
,150,000
105,000
31,200
,021,500
153,300
,325,000
^086, 000
27.200
655.500
682.700
,013,000
108,600
201,300
,655,900
,121,600
PER CENT
51.78
0.68
0.20
6.62
0.98
15.10
78.36
0.17
1.25
1.12
13.25
2,65
1.32
17.22
100.00
92
• MONSANTO RESEARCH CORPORATION •
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40i
BREAK-EVEN POINT
30-
CO
*
5
z
O
20-
Z
O
10-
PROFIT
0.8 0.6 0.4 0.2
COST
MILLS PER KILOWATT HOUR
0.2 0.4 0.6 0.8 1.0 1.2 1.4
1.6 1.8
2.0
1.0
1.0 2.0
$ PER TON OF COAL
3.0
4.0
Figure 22. Effect of Product Credit on Operating
Cost of Kiyoura-T.I.T. Process (New
Plant)
93
• MONSANTO RESEARCH CORPORATION •
-------
Table 15
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year 8100*4 Capacity
Category: Low Temperature Effluent
Name of Process Reinluft Process Flue Gas Rate
2.5
MMSCFM
MW HOD
Fixed Capital Cost $28
,600,000
ITEM TOTAL $
1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
21.
22.
23.
24.
Raw Materials & Chemicals 1
Direct Labor
Supervision
Maintenance, 5* of Fixed Capital 1
Supplies, 15$ of Maintenance
Utilities 2
Other
TOTAL DIRECT COST 8
Payroll Burden, 20* of 2 & 3
Plant Overhead, 50$ of 2, 3,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST j^
Depreciation, 10 % Fixed
CapItaT/Yr . 2
Taxes, 2$ of Fixed Capital
Insurance, 1* of Fixed Capital
Other
TOTAL FIXED COST J
TOTAL OPERATING COST 1J
COST: $/Ton of Coal 2.95
Mill/kwh 1-^
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mlll/kwh
,687,000
315,000
31,200
.430,000
211,500
,070,200
.717,900
6qf^oo
QQS.IOO
,061,700
,860,000
572.000
286.000
.718.000
.530.600
PER
31
2
0
10
1
15
61
0
7
7
21
1
2
27
CENT
.61
.33
.23
.56
.59
.30
.65
.51
.36
.87
.11
.23
.11
.18
94
• MONSANTO RESEARCH CORPORATION •
-------
40
30
O
u
u
at
D
O
O
O
Z
O
10
PROFIT
COST
MILLS PER KILOWATT HOUR
0.8 0.6 0.4 0.2
I
0.2 0.4 0.6 0.8 1.0 1.A 1.4 1.6
2.0
1.0
0 1.0 2.0
$ PER TON OF COAL
3.0
4.0
Figure 23- Effect of Product Credit on Operating Cost
of Reinluft Process
95
• MONSANTO RESEARCH CORPORATION •
-------
Table 16
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year 8100 % Capacity
Category:
Low Temperature Effluent
Name of Process Sulfacid Process piue Gas Rate
MMSCFM
1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
21.
22.
23-
2*4.
MW 1,100
Fixed Capital Cost
ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 535 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, in $ Fixed
CapftaT/Yr.
Taxes, 2% of Fixed Capital
Insurance, 1$ of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 5'29
Mlll/kwh 2,5-9
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh
61,750,000
TOTAL $
378,000
105,000
31,200
3,087,500
163,200
10,356,600
11,121,500
27,300
1,813,500
1,870,800
6,175,000
1,235,000
617,500
8.027.500
?1 ^iq.flno
PER CENT
1.56
0.13
0.12
12.69
1.91
12.59
59.30
U.ll
7.59
7.70
25.39
5.08
2.53
33.00
100.00
96
• MONSANTO RESEARCH CORPORATION •
-------
60
Q 50
U
U
Of
2 40
IS
O 30'
u 20-
at
10-
PROFIT
0.8 0.6 0.4 0.2
•I '
BREAK-EVEN POINT
COST
MILLS PER KILOWATT HOUR
0.2 0.4 0.6 0.8 1.0 1.2 1.4 t.6 1.8 2.0
t i i I i i i i i i
2.0
i
1.0
1.0 2-0
$ PER TON OF COAL
3.0
4.0
5.0
Figure 24. Effect of Product Credit on Operating Cost
of Reinluft Process
97
• MONSANTO RESEARCH CORPORATION •
-------
Table 17
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year 6100.5*. Capacity
Category: Low Temperature Effluent
Name of Process Mitsubishi Process,^ Qas R&te 2>5 MMSCFM
MW 1,400
Fixed Capital Cost$20 ,790,000
ITEM
1. Raw Materials & Chemicals
2. Direct Labor
3- Supervision
4. Maintenance, 5% of Fixed Capital
5. Supplies, 15% of Maintenance
6. Utilities
7. Other
8. TOTAL DIRECT COST
9. Payroll Burden, 20$ of 2 & 3
10. Plant Overhead, 50$ of 2, 3,
4 and 5
11. Pack & Ship
12. Waste Disposal
13. Other
14. TOTAL INDIRECT COST
15. Depreciation, 10 % Fixed
Capltal/Yr.
16. Taxes, 2% of Fixed Capital
17. Insurance, 1$ of Fixed Capital
18. Other
19. TOTAL FIXED COST
20. TOTAL OPERATING COST
21. COST: $/Ton of Coal 3'58
Mill/kwh 1.48
22. BY-PRODUCT CREDIT
23- ADJUSTED OPERATING COST
24. ADJUSTED COST: $/Ton of
Coal
Mlll/kwh
TOTAL $
10,765,000
315,000
31,200
1,039,500
156,000
600,000
12,906,700
69,300
770,900
840,200
2,079,000
415,800
208,000
2,702,800
16,449,700
PER CENT
65.46
1.91
0.18
6.32
0.94
3.65
78.46
bT4~2
4.68
5.10
12.65
2.53
1.26
16.44
100.00
98
• MONSANTO RESEARCH CORPORATION •
-------
40 -i
30 -
Z
o
o
z
o
20-
10-
BREAK-EVEN POINT
PROFIT
0.8 0.6 0.4 0.2
COST
MILLS PER KILOWATT HOUR
0.2 0.4 0.6 0.8 1.0 1.2 1.4\1.6
2.0
1.0
0 1.0 2.0
$ PER TON OF COAL
3.0
4.0
Figure 25. Effect of Product Credit on Operating Cost
of Mitsubishi Process
99
• MONSANTO RESEARCH CORPORATION •
-------
Table 18
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year e100% Capacity
Category: Low Temperature Effluent
Name of Process T.V.A-Sulfuric piue Gas Rate
2.5
MMSCFM
Acid Process
MW 1,100
Fixed Capital Cost
$56,000,000
1.
2.
3.
H.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15.
16.
17.
18.
19.
20.
21.
22.
23.
24.
ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5? of Fixed Capital
Supplies, 151 of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20* of 2 & 3
Plant Overhead, 50$ of 2, 3,
4 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, m % Fixed
CapitaT/Yr .
Taxes, 2% of £ixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL. FIXED CQST
TOTAL OPERATi*JG COST
COST: $/Ton of Coal 1-69
Mill/kwh 1.94
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh
TOTAL $
59^,000
105,090
31,200
2,800,000
420,000
8,600,000
12,550,200
27,300
1,678,100
1,705,100
5,600,000
1,120,000
560,000
7,280,000
21.535.600
PER CENT
2.76
0.48
0.11
13.01
1.95 '
39.91
58.28
.7.80
7.92
26.00
5.20
2.60
33.80
100.00
100
• MONSANTO RESEARCH CORPORATION •
-------
60
50-
«
u
« 40H
3
ts>
Z
O
>-
at
tu
a.
u
5.0
$ PER TON OF COAL
Figure 26. Effect of Product Credit on Operating Cost
of T\V.A. — Direct Acid Process
101
• MONSANTO RESEARCH CORPORATION •
-------
J. t J. J7
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year 6 10* Capacity
Category: Existing Power Plant (Low Temperature Effluent)
Callery Chemical
Name of Process Company Process piue Gas Rate
2.5
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 1"00
Fixed Capital Costl5
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,
4 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed
Capital/Yr.
Taxes, 2% of Fixed Capital
Insurance, 1JE of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 1>03
Mill/kwh 0.43
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh
,300,000
TOTAL $
1,015,000
105,000
31,200
765,000
114,750
169,000
2,199,950
27,200
508,000
535,200
1,530,000
306,000
153,000
1,989,000
4,724,150
PER CENT .
21.49
2.22
0.67
16,19
2.43
3-57
46.57
0.57
10.75
11. id
32.39
6.48
3.24
. 42.11
102
• MONSANTO RESEARCH CORPORATION •
-------
< 20
u
o
c
o
CD
O.
OJ
O
10
Break Even Point
Profit
Mills per Kilowatt
0.6 0.5 0.4 0.3 0.2 0.1
Cost
Hour
0.1 0.2 0.3
0.5 0.6
i i.
1.5
1.0
0.'5 0
$ per Ton of Coal
0.5
1.0
Figure 27. Effect of Product Credit on Operating Cost
of Gallery Process
103
• MONSANTO RESEARCH CORPORATION •
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REFERENCES
1. A. Magnus, "Composition de Compositio - Theatricum Chemicum,"
Argentarati, 4^ 929 (1613).
2. P. Phillips, Brit. Pat. 6096 (1831).
3- R. Dossie, "The Elaboratory Laid Open," London, p. 44 (1758).
4. W. Kleese, "Possible Applications of Rhenium Metal and Its
Alloys," Metalloy 5, 155 (1951).
5. B. G. Mandelik, "Conversion of Sulfur Dioxide with Low Ignition
Catalysts," U. S. Pat. 3,282,645 (1966).
6. D. Bienstock, et al., Process Development in Removing Sulfur
Dioxide from Hot Flue Gases, U. S. Bureau of Mines Report of
Investigation 5735, pp. 25-27 (1961).
7. H. Wolf, W. Goesele and G. Schachenmeier. "Removal of Sulfur
Dioxide from Flue Gases," Ger. Pat. 1,204,770 (1963).
8. D. R. Coughanowr and F. E. Krause, "The Reaction of S02 and
02 in Aqueous Solution of MnSO^/'Ind. Eng. Chem. Fundamentals
4_, 61 (1965).
9. H. Buff and A. W. Hofmann, Liebig's Ann. 113. 129 (i860).
10. G. Rienacker, Z. Electrochem. 46_, 369 U940).
11. J. Eckell, Z. Electrochem. 39., 433 (1933).
12. H. E. Farnsworth and R. F. Woodcock, Adv. Catalysis 9., 123
(1957).
13. H. M. C. Sosnovsky, J. Phys. Chem. Solids 1£, 304 (1959).
14. M. J. Duell and A. J. B. Robertson, Trans. Farad. Soc. 57,
1116 (1954).
15. E. M. Loebl, Trans. N. Y. Acad. Sci. 2_3, 491 (1961).
16. A. Sherman and H. Eyring, J. Am. Chem. Sco. 54_, 2661 (1932).
17. G. Okamoto, J. Horiuti and K. Kirota, Sci. Papers, Inst. Phys.
Chem. Res. Tokyo 2£, 223 (1936).
Preceding page blank
105
MONSANTO RESEARCH CORPORATION
-------
18. A. P. Thompson, Trans, A.I.Ch.E. 27., 264 (1931); Chem. Met.
Eng. 38., 705 (1931).
19. B. Neumann, Z. Elektrochem., 34_, 696 (1928); Z. Elektrochem.
35, ^2 (1929).
20. J. S. Streicher, Chem. Met. Eng. 37., 501 (1930).
21. T. R. Harney, Chem. Met. Eng. 37., 374 (1930).
22. G. K. Boreshkov, Trudur Pervoi Vsesoyuzroi Konferentzii po
Sernoi Krislote i sere, Soznannoi Gipokhimom NKTP, USSR,
Lenengrad, 1934, p. 106 see: CA 30 3172-9.
23. R. E. Kirk and D. F. Othmer, "Encyclopedia of Chemical Technology,"
Interscience Encyclopedia, Inc., New York, 1954, Vol. 13, p.
487.
24. A. M. Fairlie, "Sulfuric Acid Manufacture," Rheinhold, New
York, 1936, Chap. XVIII.
25. W. W. Duecker, and J. R. West (eds.), "The Manufacture of
Sulfuric Acid," Rheinhold, New York, 1959, Chap. 12.
26. Catalysis Vol. VII, ed. Paul H. Emmett, pp. 322-324, Reinhold,
New York, I960.
27. D. H. Napier, M. H. Stone, "Catalytic Oxidation of Sulphur
Dioxide at Low Concentrations," J. Appl. Chem. £, December
1958, p. 781.
28. P. Mars, and J. G. H. Maessen, "The Mechanism of the Oxidation
of Suphur Dioxide on Potassium - Vanadium Oxide Catalysts,"
Proc. Intern. Congr. Catalysis, 3d, Amsterdam 1964, 1, 266.
29. P. Mars and J. G. H. Maessen, "The Mechanism and the Kinetics
of Sulfur Dioxide Oxidation on Catalysts Containing Vanadium
and Alkali Oxides," Journal of Catalysis, 10, 1-12 (1968).
30. M. Goldman, L. N. Canjar, and R. B. Beckman, "Kinetics of
Catalytic Oxidation for Comparison of Fixed and Fluid Bed
Reactors," J. Appl. Chem. 7, May 1957, p. 274.
31. R. B. Eklund, "The Rate of Oxidation of Sulfur Dioxide with
a Commercial Vanadium Catalyst," Almquist and Wiksells,
Uppsala 1956.
106
• MONSANTO RESEARCH CORPORATION •
-------
32. W. K. Lewis and E. D. Ries, "Influence of Reaction Rate on
Operating Conditions in Contact Sulfuric Acid Manufacture,"
Ind. Eng. Chem. 17, 593-8 (1925).
33. W. K. Lewis, and E. D. Ries, "Influence of Reaction Rate
on Operating Conditions in Contact Sulfuric Acid Manufacture.
II," Ind. Eng. Chem. 1£, 830-7 (1927).
34. P. H. Calderbank, "The Mechanism of the Catalytic Oxidation
of Sulphur Dioxide with a Commercial Vanadium Catalyst: A
Kinetic Study," J. Appl. Chem. 2_, August 1952, p. 482.
35. H. Kubota, S. Mankoong, T. Akehata, and M. Snindo, "Optimum
Process Conditions for a Completely Mixed Multistage Reactor,"
Can. J. of Chem. Eng., April 1961, p. 64.
36. G. K. Boreskov, and V. P. Pligunov, J. Appl. Chem. USSR
(Zh. Priklad, Khlm. ) 6_, 785 (1933).
37. G. K. Boreskov, Chem. Abstr. 2£, 6371 (1935).
38. G. K. Boreskov, and F. J. Sokolova, J. Chem. Inc. USSR 14,
1241 (1937).
39. K. J. Brodovich, J. Appl. Chem. USSR L4 894 (1941); Chem.
Abstr. 39., 3997 (1945).
40. G. K. Boreskov, L. G. Ritter, and E. J. Volkova, Chem. Abstr.
4_3, 8251 (1949); J. Appl. Chem. USSR 2£, 250 (1949).
41. M. S. Zakarevski, and Man-Cheng Chang, Chem. Abstr... 53, 13755
(1959); Man-Cheng Chang, c.s., Chem. Abstr. 5J_, 16127 (1962).
42. P. B. Rzaev, V. A. Roiter, and G. P. Korneichuk, Urk. Khim.
Zh. 2£, 161 (I960).
43. B. Davidson and G. Thodes, A.I.Ch.E. J. 10, 568 (1964).
44. G. K. Boreskov, R. A. Tuyanov, and A. A. Ivanov, Kinetika i
Kataliz 8, 153 (1967).
45. 0. A. Hougen and K. M. Watson, Ind. and Eng. Chem., May, 529
(1943).
46. E. L. Krichevskaya, J. Phys. Chem. (USSR) 21_, 287 (1947).
47. G. H. Tandy, J. Appl. Chem. £, February, 68 (1956).
107
• MONSANTO RESEARCH CORPORATION •
-------
48. P. H. Calderbank, Chem. Eng. ProgreifBy 4£, #11, 585 (1953).
49. W. K. Lewis and E. D. Ries, Ind. Eng. Chem., 1£, 830 (1927).
50. K. Knietsch, Ber. , 31, 4093 (1901).
51. M. Bodenstein and G. C. Fink, Z. Phys. Chem., 60, 1 (1907).
52. V. A. Roiter, et al., Kinet. Katal., I, 408 (I960).
53. 0. A. Uyehara and K. M. Watson, Ind. Eng. Chem., 35, 541
(1913). ~~
5^. G. Rienacker, Chem. Tech. (Berlin) 11., 246 (1959).
55- I- E. Adadurov, M. V. Apansenko, L. M. Orlova, and A. I.
Ryabchenko, J. Appl. Chem. (USSR) 7_, 1355 (1934).
56. D. V. Gernet, A. Khitum, J. Appl. Chem. (USSR) 598 (1935).
57- G. M. Schwab, E. Kaldls, Naturewissenschaften 50, 516 (1963).
58. D. V. Gernet and A. V. Shiryueva, Phys. Chem. (USSR) 7, 450
(1936).
59. J. Scheve and E. Scheve, Z. Anorg, Allgem. Chem. 333, 143
(1964). . .
60. N. P. Kurin and N. T. Rudyuk, Izvest Tomsk Politckh Inst.
83, 163 (1956).
6l. N. P. Kurin, N. P. Fywroskaya, Izvest Tomsk Politckh. Inst.
£3, 210 (1956).
62. S. Robson, British Pat. 320,930, July 25, 1928.
63. B. Newmann, G. Heintke, Z. Electrochem. 4_3, 246 (1937).
64. N. P. Kurin, et al., Invest. Tomskogo Ind. Inst. 6jO,. No. 3 7_i
(1940).
65. Khim, Referat Zhur. No. 9, 61. (1940).
66. G. K. Boreskov and T. I. Sokolova, J. Phys. Chem. (USSR)
18, 87 (1944).
67. C. Garbato, Italian Pat. 435,452, May 18, 1948.
108
• MONSANTO RESEARCH CORPORATION •
-------
68. M. E. Pozin, I. P. Mukhlenov, L. S. Vasilesku, and L. A.
Sarkits, Zhur.Priklad. Khim. Z8, 681 (1955).
69. T. Blickle, J. de Jonge, Z. Ferenczy, and P. Kaldy, Veszpremi
Vegyip. Egyet., Kozlemen. 5., 109 (1961).
70. I. M. Eselev, I. P. Mukhlenov, and D. G. Traber, Zh. Priklad.
Khim., 37., 722 (1964).
71. ibid., 31, 972 (1964).
72. ibid., 31, 1204 (1964).
73. A. G. Siemens-Schuckertwerke, British Pat. 1,003,419, Sept.
2, 1965.
74. N. V. Kozhernikova and D. G. Traber, Zh. Priklad. Khim., 39,
1272 (1966).
75. I. G. LesoKhim, E. S. Rungantseva, and T. P. Bondarchuk, Khim.
Prom. ^2, 445 (1966).
76. Fives-Penhoet., French Pat. 1,450,441, August 26, 1966.
77. H. H. Krause, A. Levy, and W. T. Reid, J. Eng. Power, 90,
38 (1968).
78. J. R. Donavan, "The Manufacture of Sulfuric Acid," edited by
W. W. Duecker and J. R. West, Reinhold Publishing Corporation,
New York, Chapter 13 (1959).
79. V. F. Postnikov, T. I. Kunin, and A. A. Ashtasheva, J. Appl.
Chem. (USSR) £, 1373 (1936).
80. I. E. Adadurov and D. V. Germet, J. Appl. Chem. (USSR) 8,
612 (1935).
81. N. V. Kozhevhikova and D. G. Traber, Zh. Priklad. Khim. 39,
1272 (I960).
82. H. Buff and A. W. Hofmann, Ann. de Chem. Phys., 113, 129 (i860)
83. M. P. von Wilde, Ber. 7_, 352 (1874).
84. G. Glockler and S. C. Lind, "The Electrochemistry of Gases
and Other Dielectrics," Chapter XIII, John Wiley and Sons,
Inc. New York, (1939).
85. S. A. Gregory, Trans. Inst. Chem. Eng. (London), 44. 329 (.1966)
109
• MONSANTO RESEARCH CORPORATION •
-------
86. S. D. Mahant, J. Indian Chem. Soc., £, 4l? (1932).
87. M. Poliakoff, Mag. Chem. Cath. Katerinoslav (1936), 26?.
88. N. Zalogin and N. .Nechaeva, J. Phys. Chem. (USSR), 4_,
832 (1933).
89. L. A. Kolodkina and N. Nechaeva, J. Phys. Chem. (USSR), 4_,:'
(1933).
90. N. Zalogin and N. Nechaeva, J. Phys. Chem. (USSR), 6_,
(1935).
91. W. R. Browne and E. E. Stone, "Sulfur Dioxide Conversion Under
Corona Discharge Catalysis," General Electric report under
Contract PH 86-65-2 for U. S. Department of Health, Education,
and Welfare (1965).
92. Herman Mohler, ed., "Chemische Reaktionen lonisierender Strahlen,"
(Radiation Chemistry), H. R. Sauerlaender & Co., Aarau and
Frankfurt am Main, 1958, p. 235.
93. Radioisotopes - Production and Development of Large-Scale Uses,
U. S. Atomic Energy Commission Report No. 1095, Washington,
May 1968.
94. G. Kornfield and E. Wegman, "The Oxidation of Sulfur Dioxide
in Ultra-Violet Light," Z. Elektrochem. , 36., 789-91* (1930),
In German.
95. D. D. Eley, H. Pines, and P. B. Weisz, eds., "Advances in
Catalysis and Related Subjects," Vol. 18, Academic Press, New
York and London, 1968, Chapter on The Effects of Ionizing Radia-
tion on Solid Catalysts by E. H. Taylor, pp. 111-248. • '
96. P. J. Brenner, "Coal Researchers are Grappling with Sulfur, '•'
Chem. Eng., Tj., (21), 116, (1967).
97. A. V. Slack, "Air Pollution: The Control of S02 from Power Stacks,
Part III, Processes for Recovering S02," Chem. Eng. 7^ (25), 188-
196 (1967).
98. D. Bienstock, et al. , "Evaluation of Dry Processes . for Removing
Sulfur Dioxide from Power Plant Flue Gases," JAPCA, 1$, ^59-64,
(1965).
99. S. Katell, "Removing Sulfur Dioxide from Flue Gases," Chem. Eng.
Prog. 62, (10), 67-73, (1966).
110
• MONSANTO RESEARCH CORPORATION •
-------
100. "New Pilot Plants Tackle S02 Pollution," Chem. & Eng. News,
July 4, 1966, p. 36.
101. R. Kiyoura, "S02 Stack Gas Gives (NHj4)2 SO^," Chem. & Eng.
News, June 27, 1966, p. 23.
102. R. Kiyoura, "Studies on the Removal of Sulfur Dioxide from Hot
Flue Gases to Prevent Air Pollution," JAPCA, 16 (9), 488-489
(1966).
103. F. Johswich, "The Present Status of Flue Gas Desulfurization,"
Combustion, October 1965-
104. "Sulfuric Acid Process Reduces Pollution," Chem. & Eng. News,
December 21, 1964.
105. W. Moeller and K. Winkler, "The Double Contact Process for
Sulfuric Acid Production," JAPCA 18, (5), 1968.
106. H. F. Johnstone, "Metallic Ions as Catalysts for the Removal
of Sulfur Dioxide from Boiler Furnace Gases," Ind. Eng. Chem.
23., 559-561 (1931).
107. R. L. Copson and J. W. Payne, "Recovery of Sulfur Dioxide as
Dilute Sulfuric Acid. Catalytic Oxidation in Water Solution,"
Ind. Eng. Chem. 25_, 909-916 (1933).
108. M. K. Grodzovskii, J. Phys. Chem. (USSR) 6_, 496 (1935).
109. J. W. Walthall, P. Miller, and M. M. Striplin, Jr., Trans. Am.
Inst. Chem. Engs. 4]., 110 (1945).
110. P. J. Brenna, "Coal Researchers are Grappling with Sulfur," Chem.
Eng. 74. (21), 114-118 (1967).
111. G. Tarbutton, J. C. Driskell, T. M. Jones, F. J. Gray, and C. M.
Smith, "Recovery of Sulfur Dioxide from Flue Gases," Ind. Eng.
Chem. _4_9, 392-395 (1957).
112. L. I. Sashtanov and V. P. Ryzhov, Izvest. Teplotekh. Inst. 7,
37 (1935).
113. A. V. Slack, "Air Pollution: The Control of S02 from Power Stacks "
Chem. Eng., 7_4. (25), 118-196 (1967).
Ill
• MONSANTO RESEARCH CORPORATION •
-------
114. S. Ludwig, "Antipollution Process Uses Absorbent to Remove
S02 from Flue Gases," Chem. Eng. 7_5 (2), 70-72 (1968).
115- D. Bienstock and J. H. Field, "Process for Removing Sulfur
Dioxide from Gases," U. S. Pat. 3,150,923, September 29, 1964.
116. D. Bienstock, J. H. Field, and J. G. Myers, "Process Development
in Removing Sulfur Dioxide from Hot Flue Gases," Bureau of Mines
Report of Investigation 5735, 1961.
117. T. T. Frankenberg, "Removal of Sulfur from Products of Combustion,"
Proc. Am. Petrol. Inst. Sect. Ill, 45_ (3), 371 (1965).
118. A. V. Slack, "Air Pollution: The Control of S02 from Power Stacks,"
Chem. Eng. Jj. (25), 188-196 (1967).
119. S. Katell, "Removal Sulfur Dioxide from Flue Gases," Chem. Eng.
Prog. 62_ (10), 67-73 (1966).
120. "Coke Cleans Flue Gas in German Process," Chem. Eng. 7jl_ (22),
94-98 (1967).
121. D. Bienstock, et al., "Evaluation of Dry Processes for Removing
Sulfur Dioxide from Power Plant Flue Gases,"'JA'PCA, 15, (10),
459-464 (1965).
122. P. J. Brennan, "Coal Researchers are Grappling with Sulfur,"
Chem. Eng. 7_4_ (21), 114-118 (1967).
123- R. P. Hangebrauck and P. W. Spaite, "Controlling the Oxides of
Sulfur," JAPCA, .18 (1), 5-8 (1968).
124. Firmenschrift der Lurgi-Appartebau GmbH, Frankfurt/M.
125. W. Broocke, "Aussichten fur eine Praktische Anwendung von
Abgasentschwe fellungsverfahren," Staub-Reinhalt. Luft,
2£ (3), 101 (1968).
126. Z. Tamura, S. Hori, and H. Jakashina, "Removal of S02 from Stack
Gas by Activated Carbon," preprint, 1967.
127. K. Emicke, "A Method of Removing Sulphur Dioxide from Gas Con-
taining Sulphur Dioxide," British Pat. 1,107,626, March 1968.
128. H. Wolf, et al., "Method of Removing Sulfur Dioxide from Flue
Gas," German Pat. 1,204,770, June 1966.
112
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129- S. Ludwig, "Antipollution Processes Uses Absorbent to Remove
S02 from Flue Gas," Chem. Eng., 75_ (2), 70-72 (1968).
130. R. Coleman, "The Outlook for Fertilizers," Chem. Eng. Prog.,
64:7, 68-71 (July 1968).
131. B. Kadlec and A. Regner, "Oxidation of Sulfur Dioxide on
Vanadium Catalyst in the Region of Internal Diffusion, I.
Theoretical Part."
132. B. Kadlec and V. Pour, "Oxidation of Sulfur Dioxide on
Vanadium Catalyst in the Region of Internal Diffusion, II.
Theoretical Part."
»
133. A. Regner and A. Simecek, "Kinetics and Mechanism of Sulfur
Dioxide Oxidation on a Vanadium Catalyst, III. Correlation
of the Data with Equations Derived for the Reduction-
Oxidation Mechanism," Collection Czech. Chem. Coramun. V.
33, 8 pp. 2388, 2526, 25^0, (August 1968).
113
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APPENDIX I
CONVERSION EFFICIENCY AND RATE EQUATION GRAPHS
Preceding page blank
115
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APPENDIX I
INDEX
Figure No. Title Page No,
1 Effect of Contact Time on Conversion at 118
500°C (calculated values for flue gas
application)
2 Effect of Contact Time on Conversion at 119
450°C
3 Effect of Contact Time on Conversion at 120
4 Effect of Contact Time on Conversion at 121
5 Effect of Contact Time on Conversion by 122
Mars & Maessen Equation at 400°C
6 Effect of Contact Time on Conversion at 123
Preceding page blank
7 Effect of Contact Time on Conversion at
375°C
8 Effect of W/F Ratio: on Conversion at 500°C 125
9 Effect of W/R Ratio on Conversion at 450°C 126
10 Effect of W/F Ratio on Conversion at 425°C 127
11 Effect of W/F Ratio on 'Conversion at 425°C 128
12 Effect of W/F Ratio on Conversion by 129
Mars & Maessen Equation at 752°F
13 Effect of W/F Ratio on Conversion at 400°C 130
14 Effect of W/F Ratio on Conversion at 375°C 131
117
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KATF EQUATIONS FROM TABLE 5:
Mars & Maessen (Catalyst #3)
——^— Mars & Maessen (Catalyst #1)
•—'—•—-—- Eklund
— — Calderbank
0.1 0.2
CONTACT TIME, »»condt
0.3
Figure 1. Effect of Contact Time on Conversion at
500°C (calculated values for flue gas
application)
118
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100-1
90
Mars & Maessen (Catalyst #3)
, Mars & Maessen (Catalyst #1)
Eklund (at H60°C)
—--— Calderbank
X
Breaches 90$
@ 0.9 sec
X
X
X
0.1 0.3
CONTACT TIM!. i«condi
0.3
Figure 2. Effect of Contact Time on Conversion at 450°C
119
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Mars & Maessen (Catalyst #1)
Goldman, et al
Mars & Maessen (Catalyst #3)
0.1 0.2
CONTACT TIME, tacendi
0.3
Figure 3. Effect of Contact Time on Conversion at
120
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Eklund
— Calderbank
Davidson-Thodos
0.1 0.2
CONTACT TIMC. t«condi
0.3
Figure 4. Effect of Contact Time on Conversion at
121
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100-
90
^Catalyst #3
Catalyst #1
0.1 0.2
CONTACT TIME, i*cendi
Figure 5. Effect of Contact Time on Conversion by Mars &
Maessen Equation at 400°C
122
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-Goldman, et al
Davidson & Thodos
3.0 4.0 3.0 4.0
CONTACT TIME, iccondt
Figure 6. Effect of Contact Time o,n Conversion at 400°C
123
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lOOi
90
Davidson-Thodos
... Goldman, et al
i.o 2.0
CONTACT TIME. ««condi
3.0
Figure 7. Effect of Contact Time on Conversion at 375°C
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100-1
90
Mars & Maessen (Catalyst #3)
Mars & Maessen (Catalyst
_—. Eklund
Calderbank
1.0 2.0
W/P, Q" i«c/meU 5O»«ID
3.0
Figure 8. Effect of W/F Ratio on Conversion at 500°C
125
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I OO-i
90
Mars & Maessen (Catalyst #3)
Mars & Maessen (Catalyst #1)
Eklund (at 460°C)
Calderbank
^Reaches 90$
% 13 x 106
1.0 2.0
W/F, gm »«c/mol« SO,FEED x!0~*
3.0
Figure 9. Effect of W/F Ratio on Conversion at
126
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Mars & Maessen (Catalyst #3)
Goldman, et al
Mars & Maessen (Catalyst #1)
i i i i
i.o a.o ).o
W/f, QM ••€/•»!• SO,»10-«
Figure 10. Effect of W/F Ratio on Conversion at
127
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100-
90
•Eklund (at
-Calderbank
.Davidson-Thodos
10.0 90.0
r, gn i«c/«eU
30.0
Figure 11. Effect of W/F Ratio on Conversion at k25°C
128
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Catalyst #3
.— Catalyst #1
2.0 3.0 4.0 .5.0
W/F. gm t«c/«eU *O,««0 "lO"4
6.0
Figure 12. Effect of W/F Ratio on Conversion by
Mars & Maessen Equation at 752°F
129
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lOO-i
Goldman, et al
Davidson-Thodos
Calderbank
10.0 20.0
W/F. g« icc/aol*
30.0
Figure 13. Effect of W/F on Conversion at 400°C
130
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lOO-i
90
Davidson-Thodos
,-jQoldman, et al
10.0 20.0
r, 9m ••c/meU SO.FMD
JO.O
Figure 14. Effect of W/F Ratio on Conversion at 375°C
131
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APPENDIX II
MATHEMATICAL MODELS AND COMPUTER PROGRAM LISTINGS
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133
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APPENDIX II
INDEX
Title , Page No.
Reactant Partial Pressures and Conversion Fraction 136
at Equilibrium for the Reaction S02+l/2 02tS03 *
Partial Pressure and Conversion Fraction Computer 139
Program Listing
Calculation of Percentage Conversion and Weight of
Catalyst in an Isothermal Catalytic Reactor as a
Function of Contact Time
Conversion Percentage and Catalyst Weight Calculation
Computer Program Listing
Calculated values
Preceding page blank
135
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REACTANT PARTIAL PRESSURES AND CONVERSION FRACTION AT
EQUILIBRIUM FOR THE REACTION SO? -I- 1/2 0? t~S07
Ref. 11 gives an approach for calculating reactant partial pressures
and the equilibrium conversion fraction for the reaction S02 + 1/2 02
-303, However, the equations for the reactant partial pressures,
= 2'0° av •
2.00 - a(l-y)
PSCH = 2.00 a(l-y) ,
PS03 2.00 - a(l-y) (2)
- 3a + ay
2.00 - a(l-y)
[a is the initial partial pressure of S02 (atm.) in the burner gas
obtained by the burning of sulfur in air, and y is the fraction of
a remaining unconverted at equilibrium] were derived for the
condition that sulfur is burned in air of 2 atm. pressure, and thus
part of the oxygen is used to oxidize the sulfur to S02 • In a
catalytic reactor for oxidation of S02 in flue gases, different
conditions prevail, the pressure is usually well below 2 atm. and
che initial partial pressures of S02 and 02 are lower than those for
the case discussed above. The following will give the equations
representative of the conditions prevailing in a catalytic reactor
for the treatment of flue gases:
Let,
PTOT = Total pressure of all flue gas components in the
catalytic reactor in atm.
a = initial vol. fraction of S02 in the flue gas
b = initial vol. fraction of 02 in the flue gas
136
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Assume that y is the fraction of a remaining unconverted at
equilibrium. Then, at equilibrium
- (b-0.5a(l-y)}PTOT
l-0.5a(l-y)
l-0.5a(l-y)
- a(l-y)PTOT
- l-0.5a(l-y)
Also, for the reaction
S02 + 1/2 02 Z SO3
v - PS03
(PS02)-(P02)'/
where K is the equilibrium constant.
Substitution of Eq's. (4), (5), (6) in Eq. (7) and rearranging
yields
0.5aUPTOT'j'K 2-l }y 3+{ i PTOT)-K 2 (b-0. 5a)-(l-l. 5a; }y 2
+{2-1.5a}-y+(-l+0.5a} = 0
Tne cubic equation ^8) can be solved for y and for a range of
values cf a, b, PTOT and K - The only y values ;roots) with
physical meaning are thcsepwhich are positive and j_l. The cal-
culated values of y can be substituted in Eq's. (lj, (2), and (3)
to determine the reactant partial pressures. The conversion
fraction at equilibrium, i.e., the fraction of S02 converted into
SO 3, as equal to 1-y.
137
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Temperature dependence of the partial pressures of the reactants
can be evaluated by introducing K as a function of temperature
into Eq. (B), Ref. -i gives the pfollowing relationship between
K and T (T = absolute temperature, CK).
log Kp =
1.956
- 1.678
(9)
A computer program was developed for the solution of Eq's, (8)
and (9) and for the computation of the partial pressures and the
conversion fraction. This program allows for a parametric
evaluation of the reaction S02 + 1/2 02* SOa for a range of total
pressures, initial vol % of S02 and 02, and temperatures.
The next two pages show a FORTRAN-IV listing and a sample print-out,
respectively.
138
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Table 1
FORTRAN-IV PROGRAM FOR CALCULATING EQUILIBRIUM
CONVERSION VALUES
// FOR
•rOCSICARD.1132 PRINTER)
•EXTENOED PRECISION
* LIST ALL
REAL KPtKPZtlNT
DIMENSION BI5I.AAI6)
READ(2t»ILtlAA< IUI-1.LI
REAO(2t9)Kt(B(I>iI-ltKI
5 FORMAT! 15 .6F5.0)
00 60 J'ltL
00 6A I-1.IC
WRITEI3.109I
109 FORMAT) •!')
PTOT-I.l
A*AA(JI
A«A*100.
WRITE! SflOaiPTOTtAiBII)
10B FOHMATllOX.tTOTAL PRESSURE ( ATM) • ,F8.2 .//t
I IOXf 'INITIAL VOLUME PERCENT OF SO2- « ,F8.2// .
? 10X. 'INITIAL VOLUME PERCENT OF 02«'tF8.?//l
AaA/100*
B(|I«B(M/100.
WRITEOilOO)
100 FORMAT (5Xt 'TEMP', 16X.'P$0?'tl8X.'P02'.lSX.»PS03'.13X.«KSUBP'.l6X.
CONi'/XI
N«0
10 N'N*1
KP«10.0»««.«)56.0/T-*.678I
0«0.9»A»TTOT»UP2-1.0) I
AR*R(||
P«lPTOT»KP?»IAR-0.»»A)-U»0-l.S»AI
P«P/0
0 = 0/0
R=H/0
TEST»0»P»0+R
IF(TEST)litl6,16
15 fel.O
fiC TO 50
Ib INT.0.1
v-l.l
30 SV«v
20 r«Sr
2«S7
INT«INT/10
IM INT-1.0F.-10I50.50.30
50 CONTINUF
PSO?-2.0*A«W:?*0-A*I1.0-VI)
P02» IO.«.2-3.0»A*A»Y)/I2.0-A»I1.0-Y)
EC»(1*0-VI*100.
MRITEI 3 t 1031 T.PS02.P02.PSO3.KP.FC
103 FORMAT(2X.E12.5.5I8X.E12.5».I»)
IF(T-1200.0)10.60.60
60 CONTINUE
CALL EXIT
END
139
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Table 2
SAMPLE PRINT-OUT FROM COMPUTER PROGRAM FOR
CALCULATING EQUILIBRIUM CONVERSION VALUES
o
z
w
>
z
o
a
m
to
m
o
i
o
o
3
T)
O
O
z
TOTAL PRESSURE (ATM) 1.10
INITIAL VOLUME PERCENT OF SO2-
INITIAL VOLUME PERCENT OF O2-
TEMP
O.SOOOOE 03
0.55000E 03
0.60000E 03
0.64OOOE 03
0.700OOE 03
0.79000E 03
0.80000C 03
0.05000E 03
0.90000E 03
0.95000E 03
0.10000E 04
0.10900E 04
0.11000E 04
0.11500E 04
0.12000E 04
PSO2
0.22271E-07
0.33979E-06
0.30520E-05
0.17291E-04
0.67437E-04
0.19978E-03
0.47394E-03
0.91191E-03
0.14395E-02
0.19275E-02
0.22963E-02
0.25414E-02
0.30000E-02
0.30000E-02
0.30OOOE-02
0.30
2.50
P02
0.20580E 00
0.20560E 00
0.20580E 00
0.20591E 00
0.20583E 00
0.20568E 00
0.20599E 00
0.20617E 00
0.20638E 00
0.20657E 00
0.20672E 00
0.20681E 00
Q.20700E 00
0.20700E 00
0.20700E 00
PS03
0.30044E-02
0.30041E-02
0.30014E-02
0.29871E-02
0.29369E-02
0.28044E-02
0.25298E-62
0.20912E-02
0.15627E-02
0.10740E-02
0.70474E-03
0.45919E-03
O.OOOOOE 00
O.OOOOOE 00
O.OOOOOE 00
KSUBP
0.17139E 06
0.21i23E 05
0.38194E 04
0.80433E 03
0.2S234E 03
0.85113E 02
0.32885E 02
0.1420VE 02
0.67401E 01
0.34!>81E 01
0.18967E 01
0.110I5E 01
0.67213E 00
0.42811E 00
0.28313E 00
EOU. CON.
0.99999E 02
0.99988E 02
0.99424E 02
0.97755E 02
0.93349E 02
0.84221E
0.69634E
0.52O52E
0.35782E
0.23483E
0.15302E 02
O.OOOOOE 00
O.OOOOOE 00
O.OOOOOE 00
02
02
02
02
02
-------
Calculation of Percentage Conversion and Weight
of Catalyst in an. Isothermal Catalytic Reactor
as a Function of Contact Time
The calculations described below will depend on the availability
of a kinetic equation for the reaction expressing the rate of
conversion, in this case of S02, as a function of the various
relevant variables.
V - f(PS02> P02> PS03' T> Kl"Kn)
where: v is the velocity of the conversion per gram of cata-
lyst (moles SC>2/sec-g. )
T is the temperature (°K)
Ki««Kn are various parameters of the kinetic equation
such as equilibrium constants which may or may not also
be functions of T,
Px is the partial pressure of component x (atm.)-
Using this expression it is possible to calculate the rate of
the conversion for a given set of reaction conditions. What
we would then like to do is to be able to use the velocities
calculated to compute the composition of the gas stream passing
through the catalyst as a function of time.
If we assume that we start out with a feed stream of initial
composition PA , Pcn , and PQn at temperature T, we can use
\J2 £>U 2 oU 3
equation (1) to calculate v0 the instantaneous rate of reaction
under the initial conditions.
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Using Equation (2) we can calculate the change in the number of
moles of S03 produced (Amcn ) in a time At. Since
AmS03 " v°At (2)
the change in the pressure of S03 (APSQ ) is directly pro-
portional 'to AmQri (Eq. 3) it should be possible to calculate
oU^
the former quantity.
PS03
Then using the expressions:
so3
Po2 •
PS02 =
PS03
Po2 -
PS02 "
PS03 "*"
1/2 A P
APS03
APS03
the concentration of the feed gas at the end of this time incre-
ment At can be calculated. This process can be repeated for the
next time increment At using the new pressures calculated in place
of the initial pressures, to calculate the composition of the gas
stream at the end of the time period 2At.
Repetition of this process will allow one to calculate the com-
position of the gas stream at any given contact time. This in-
formation can be substituted into Equations (7) and (8) to calcu-
late fraction of SC>2 converted (a) and percent conversion (fa) of
SC>2 for that time period.
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PS03
p - (7)
*S02
P
S02
S03
%a = 5 - x 100 (8)
r
All that remains in the way of making these calculations is
being able to find a value for Kc,the proportionality factor
between APQn and Am^ . This K can be derived as follows:
oU i><-' C
Let F be the volume flow rate of gas stream in
the reactor at temperature T. If T is constant
then the volume flow rate, F, is a constant through-
out the reactor. Therefore, the incremental
time, At, it takes for a unit volume of gas to
flow through an incremental volume, AVC, of
catalyst bed is given by:
(9)
For a catalyst of bulk density, p, AVC can be related to the
incremental weight of the catalyst AW as follows:
AVC » AW/p (10)
Substituting AVC in Equation (9) yields
or
W = pFAt (12)
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UsJ.ng the ideal gas law, the number of moles of S03 passing
through the catalyst bed at any contact time t per gram of
catalyst is given by:
PS03F
mS03 ~ RTW
Therefore,'during a time increment at
A "D f
or-1.-. r
:>03
At
RTW
If we substitute Equation (12) into Equation (14) we obtain
APSO/
or
(16)
If we compare Equation (16) with Equation (3) we find that
Kc = pRT (17)
A further quantity that would be helpful to calculate is the
total amount of catalyst necessary to effect a particular per-
cent conversion of S02- This quantity can be calculated in
the following manner. If we assume that as the gas passes
through an increment of catalyst weighing dW the change in
fraction converted is da. Then:
dW = -| da (18)
where v reaction rate is a function of a (conversion)
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Since we do not have the explicit form of the expression for
the instantaneous velocity v as a function of a at constant
temperature, we cannot integrate Equation (18) analytically.
We can, however, integrate this equation numerically using
the difference formula:
yk
where W is the weight of catalyst which
has contacted the gas stream after
the kth increment of contact time.
if
V = The velocity of reaction at
the start of the kth incre-
ment of contact time.
V
Aa = The change in fraction S02
converted during the kth
increment of contact time.
We have written two programs for an IBM 1130 computer which makes
the calculations described above. The mainline program CONV.
handles all the input and output. In addition, it calculates the
pressures of reactants and products in the gas stream, conversions
and the W/P ratio as a function of contact time. This program is
listed on pages 149 through 152.
CONV. contains a call to the subroutine REACV. which calculates
the velocity of the conversion in moles of 803 per second per
gram of catalyst as a function of partial pressures of the com-
ponents and the absolute temperature.
145
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An illustration of one form of the subroutine REACV
which calculates the velocity of reaction according to the
equations of Davidson and Thodos is listed on page 76.
It is assumed in all these calculations that the conversion reac-
tion takes place isothermally . We have also assumed that the
bulk density of the catalyst bed is 0.75 g/cc.
The input and outputs from these programs are arranged as
follows :
INPUT
Card 1
Case title
Columns 1-80
Card 2
Til
TP
TFV
NC
integration interval (sec.)
printing interval (sec.)
upper limit of integration (sec.);
lower limit is zero
number of equation constants to be
read in on Card 4
> unformatted*
Card 3
PO
P2
P3
initial pressure of 02 in gas stream
(atm.) I
initial pressure of S02 in gas stream
(atm.) > unformatted
initial pressure of SO in gas stream I
(atm;) 3 J
"unformatted input must be separated on data cards at least
one blank. The subroutine READP which allows unformatted
input can be furnished on request.
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Card 1
C(l) » temperature (°C)
C(2)
unformatted
equation constants
C(NC)
a typical .Input listing is given on page 76.
OUTPUT
The following information is printed by the computer as output
1) case title
2) integration interval (sec.)
3) print Interval (sec.)
4) K(EQ) the equilibrium constant for the 02-S02-S03
equilibrium
5) temperature in degrees centigrade
6) P02(F) the initial feed composition of 02 (atm.)
7) PSQ (F) the initial feed composition of S02 .(atm.)
8) PSQ (F) the initial feed composition of S03 (atm.)
9) C(2-10) the values of C(2)«"C(NC)
Then at each value of contact time with corresponds to a print
interval are printed out.
1) contact time (sec.)
2) P02 the 02 composition in the gas stream (atm.)
3) Pon the SO composition in the gas stream (atm.)
2
4) P _ the S03 composition in the gas stream (atm.)
oU 3
5) PCONV the percent conversion
6) W/F the weight of catalyst used up to that time
divided by the molar feed rate (g-sec/mole S02)
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The operator has several options for running this program. If
data switch 1 is off then the computer will start a new case as
soon as it detects that the contact time for which calculations
are being made is equal to or greater than TPV. If data switch
1 is on then the computer starts polling the other'data switches
If they are all off, it will continue to wait until one of them
is turned on or data switch 1 is turned off.
If data switch 2 is on then the computer runs the same case
again at a new temperature. The data on the new temperature
is laid out as on Card 4.
If data switch 3 is on the old case is rerun with a new TFV(n)
[the value of TFV after the old case has been run n times under
the control of data switch 3] given by:
TFV(n) = TPV(n-l) + TFV(l)
If both data switches 2 and 3 are on at the same time, switch
3 will be ignored until data switch 2 is turned off.
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C PROGRAM CONV - CALULATES THE PERCENT CONVERSION AND WEIGHT OF
C CATALYST IN A CATALYTIC RFACTOR AS A FUNCTION OF CONTACT TIMF.t
C TEMPERATURE AMD INITIAL FECD STKKAM COMPOSITION
C _ _
c ..... RniHUL'irnFNSTTvor rAr«tV5T-Bcn- 7fT.yc;c".T
C Clll • TEMPERATURE ( OEG. C.)
C C ( 2-101- MISCELLANEOUS CONSTANTS USFO IN THE CALCULATIONS
C F.KP- EQUILIBRIUM CONSTANT FOR 02- S02- SOI EQUILIBRIUM.
C FCONV FRACTION OF S02 CONVERTED
C NC« THE NUMBER OF EQUATION CONSTANTS TO BE READ IN _
C PO- INITIAL PARTIAL PRESSURE OF OXYGEN (ATM.)
C P2«lNlTlAL PARTIAL PRF.SSURE OF S02 (ATM. I
C P3-INITIAL PARTIAL PRESSURE OF 503 (ATM. I
C P02- INSTANTANEOUS PRESSURE OF 02 I ATM. I
C PS02=INSTANTANEOU5 PRESSURE OF S02 IATM. ) ___
T ~ >»50j«IN5lfl1IA,NtUWi" PKtSiUKt OF "SOTTJTTMil
C R-GAS CONSTANT (CAL./ (MOLE-DEG. K.tl
T-CONTACT TIME (SEC.) .
C TFV-FINAL VALUE OF CONTACT TIME. (SEC.)
C TI I» INTEGRATION INTERVAL (SEC. I
C TK« TEMPFRATURE (CEG. K.I
'C ..... T- TPiPRINT INTEWAT" C'SFCi J .......... ". ------------------------
C V» VELOCITY OF REACTION IMQL. £03'(SEC.-G. CATALYST ) I
C M* WEIGHT OF CATALYST. CONTACTED/MOLAR PLOW RATF. (G. -SEC. /MOLE S03 1
C .
DIMENSION TITLE(20)t2<20).IN(S>
COMMON C( 101 •V.P02.PS02«PS03
Listing of Computer Program CONV
149
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PS03-ZI3)
PO-P02
P2«PS02
P3-PS03
C READ IN ANY CONSTANTS USED IN THE CALCULATION
C
140 CALL READP(NC(C(INI
TC-C(l)
TK-TC+273.16
150 WRITE!3.1002 IT 11.TPI.EKP.TC.P02.PS02.PS03
WRITE(3tl003)ELP/2.
IFIL-1 1181 .183. 1B3
1P1 FC1-FCONV
L = l
r,0 TO 106
183 FC2-FCONV
FCI-FCONV
1«6 IFITP-Tl 190.190.170
190 PCONV»rCONV»100>
C
c PRINT OUTPUT
WR I TE I 3 • 1005 ) T ,P02 .PS02 .PSO3f PCONV.W, V
1FIPCONV-99. 1191.195.195
191 IFITFV-T 1195,195.160
C
C OATSW 1 OFF- GO TO NEW CASF.. ON-POLL OTHER DATA SWITCHES
195 CALL DATStfUtJTAt
IFUTA-1 1200.200(100
C
C OATSW 2 OFF- CONTINUE . ON-RUN OLD CASE WITH NEW TEMPERATURE
C
IFUTB-1 1210(204(210
204 T«0.
TP«0.
V-0*
W-0.
~^ "~" r c i • o 9 " ~ ..... *™"~ — — "•"""
(Continued)
150
• MONSANTO RESEARCH CORPORATION
-------
PS03>2i3l
PO»P02
P2-PS02
P3-PS03
C READ IN ANY CONSTANTS USED IN THE CALCULATION
C .....
140 CALL READPINCtCtIN)
TC-CI1)
TK-TC*273.16
1*0 WR1TE(3>1002)TH«TPI tEKP»TC»P02»PSb2tP503
WRITE(3.1003MC«I).I-2.NC)
WRITEI3tl004)
160 TP»TP*TP.I
170 T-T+TII
C SUBROUTINE REACV- CALCULATES VELOCITY OF REACTION
C
ISO CALL REACV
C '
C CALCULATE CONVERSION NEW FEED GAS COMPOSITION AND X
0»V»TK»82.05»BD
DELP=0«TII
PS03«PS03»DELP
PS02«PS02-DELP
P02-P02-DELP/2.
IF(L-l)iaiil83fl63
Ifll FCl'FCONV
L-l
00 TO 186
1H3 FC2<>FCONV
199 •<»W*(FC?"fCH»l.W
FC1-FCONV
1A6 lF(TP-TU90tl90.170
190 PCONV«FCONV»100.
C
C PRINT OUTPUT
WR I TE I 3 ,1005 ) T tP02 ,PS02 tPSOS tPCONV.W, V
IF(PCONV-99.)191,195»194
191 IF(TFV-T»195. 199*160
C ...
C _ OATSW 1 OFF- GO TO NEW CASEt ON-POLL OTHER DATA SWITCHES
195 CALL OATSWIltjTAI
IF«JTA-ll200.?OOtlOO
C
C DATSW 2 OFF- CONTINUE • ON-RUN OLD CASE WITH NEW TEMPERATURE
C
IFIJTB-1 1210,204.210
204 f =0.
TP-0.
V-0»
w-o,
(Continued)
151
• MONSANTO RESEARCH CORPORATION
-------
L«0
TFV'TFVI
P02-PO
PS02-P2
WRITE (3. 1007)
GO TO 1*0
C .
C fiATSW 3 OFF-CONTINUEt ON-KUN OLD CASE CONTINUING THt INTLGRATION
C FROM N«TFV TO (N*1)»TFV. ' .
CALL DArSWI3tjTC)
IF(JTC-lll95i220.19b
220 TFV«rFV*TFVl
TiO TO 160
1000 FORMAT (?OA<»)
lU'Ji rrjKIAI I irfT t20A*//l
1002 FORMATOXt'lNTEGRATION INTERVAL" • tF8.4 i6X, 'PRINT INTERVAL- •
'K(EO)'» »iE12.5.6X.»TEMP.(C.)« • »F9.2// JX. • P02(F )= '
«Pi02lF)« «tE12.5,3Xi«VS03(FI« '.EU.5//1
1003 FORMAT(2Xi'Ct2-lOI« ' t9( El 2.S t2X )// I
100*. FORMAT(lX,lT!ME(SEC)lf6X,«P02tilC
-------
"" ~5TJHRUUTTNE"RFACV" "~" "" ' .
DAVIDSON AND THODOS A.I.CH.E.J. VOL.10 568 (1964)
COMMON C(10).V.POtP2tP3»EKPtTK
IF(C(2)-0.»20»10t20
1O T K — T V 4fc 1 O
vJ I P*» — I lx~ 1 • fj
PN=1.-PO-P2-P3
C(3)=PXP(50400./(R#TlC)-39,<»2/RJ
C ( 5 ) =EXP ( -444 00 . / ( R*TK) -»-44 • 1 3 /R )
C(6)=FXP(-36350./(R»TK)+29.77/RI
I H^F
V2=( l.+C(3)*P2-t-C(A)«PO»*.b+C(5)*P3*C(6)«PN)*«2.
V=C(2)*P2*PO*#.5*V1/V?
V=V/3600.
RFTURN :
FND
Listing of Typical Subroutine REACV.
FOUATIONS OF DAVIDSON AND THODOS
0.001 0.1 1.0 2
0.028 0.003 0.0
-------
APPENDIX III
CATALYST DATA SHEETS
Preceding page blank
155
• MONSANTO RESEARCH CORPORATION
-------
APPENDIX III
INDEX
Catalyst Phase Catalyst Base Page No.
Solid Vanadium 158
Iron 186
Chromium 191
Platinum 191*
Carbon 205
Manganese 208
Miscellaneous 209
Liquid Iron 213
Manganese 214
Miscellaneous 216
Gas Miscellaneous 217
Preceding page blank
157
MONSANTO RESEARCH CORPORATION
-------
CATALYST - VANADIUM
1
3
U
6
7
B
CitilyM
VJ«;
Ja°5
Vanadium
V3°5
"
vaoc
Vanadium
Type of
Catalysis
solid-
Gas
-
II
11
»
"
Temp.
300 °C
360 "c
U30-
e;3n°r
365 °C
500 °C
Weight %
Promoter
»
Potash
Ha n_K-fy^
*
"
Weight %
Support
Material
Diatom-
Anemia
earth
cont«.
5-20J6
A10°1
E 3
Weight %
Catalyat
tr
«• 5
ii
Porosity
Surface
Aree
Gas
Flow
Rate
cc-cat-
ftlyst
Contact
Time
n
n
so,
Cone.
11
II
0.35*
°2
Cone.
20.2*
19- 7*
IQ.CM:
Convereion
Efficiency
s*.
z
o
•o
m
W
2
•o
o
z
o
o
•o
TJ
O
a
5
z
Ul
-------
CATALYST - VANADIUM
10
11
12
13
1«
15
Catalyst
*.'/
Vanadium
Vanadium
Vanadium
V2°5
"
Vanadium
V3°5
w
Type of
Catalysis
Gas
11
"
"
^
n
Temp.
550 °C
iiso'c
«50"C
iteo'c
«70°C
soo'r:
390-C
5oo°c
420°C
'125 °C
Weight %
Promoter
TOH-K,0
KHS01(-
lili*
Weight %
Support
Materiel
N.Belaks-
JiAa;
Weight %
Catalyst
v,,o.-8*
~ -* •
.-
•7
Porosity
O.6"5cc,<
ii
"
Surface
Area
Gas
Flow
Rate
nVm'
loui-ly
ipace
vel=140<
Contact
Time
SO,
Cone.
7*
n
n
oa
Cone.
11*
n
f*
Conversion
Efficiency
86.3*
97.8*
97-ljS
96; I*
9«.»
22*
30.5*^
50*
Reference
»
CA-"i9-7Ol<}r
rA-1Q-P-7Ul1i
b
CA-60-2369a
CA-60-2^69e
CA-6o-2370d
CA-60-23VOI'
Remarks
Fe-jO, on 'V catalysts were
studied.
Polish catalvat was onmnafg^J
with other 6 important catalysts
NO cnan^e in catalyst activity
occurred after a 5-hr treat-
nent at 700°C.
POTO rArtM u» a HO-HOnA . Mpthnd
3f preparation of catalyst Is
ilfio renorted.
Promo tine action of LI. Rb.
Ce, La, Ce, Fr and Hd were
The activltv'of series of
catalysts were studied.
Optimum shape and eraln plze
Sffecta of alkali metal
iulfates on the activity of V,Cr
}ared of these cmoda- and thejlr
ictivltv and structure s£i{A4e^'
o
z
M
Z
5
m
>
X
n
i
o
o
X
•o
o
X
O
z
-------
CATALYST - VANADIUM
IB
19
20
22.
Catalyst
V3°5
""
Vanadium
Vanadium
vc°-
-f •
Tvpeof
Catalysis
Qaa
"
"
"
"
Temp.
°C
i»oo-55C
•c
500 °C
-
>1)00'C
Weight %
PrOHKIlVI
Weight %
Support
Material
Fiinod
infusor-
ial earth
Weight %
Catalyst
to fL?*
Porosity
Surface
Area
Gas
Flow
Rate
Contact
Time
SO,
Cone.
10*
7.5*
o,
Cone.
-
Conversion
Efficiency
88.68*
Reference
'•
CA-6O>-3527d
CA-60-H81l3f
CA-60-15t68c
CA-6l-10336b
Remarks
:atalyst was studied. The
legree of poisoning of catalyst
>v SiFi. was determined.
Sffect of K-0 on the aetivitv
if V-catalysts were deter-
olned. Temp- at which catalyst
iddn. of rUO.
.
Series of catalysts were
prepared to find an active
catalvst with a suitable DO re
structure.
Catalytic oxidation of SOn in
a fluldlzed bed. With an" In-
crease of the SO^ concn. from
7.S* the decree of oxldri. of
it to SO-, is reduced
neKllelblr.
carriers was independent of
structure and shape, and the.
activity depended on' the
resulting surface" of catalyst.
A table allows the reactor to '
be adjusted to changes of
throughput and catalyst
activity.
.
o
(ft
z
o
m
(A
>
o
o
H
O
-------
CATALYST - VANADIUM
24
25
2b
27
28
29
Catalyst
Vanadium
Vanadium
Vanadium
Vanadium
Vanadium
Vanadium
;V,O,-K-
vanaAaK>1
Type of
Catalysis
r.as
ii
ff
Ji
aolld-
JTA.Q
11
"
Temp.
542-63
°r
420°C
480-
finov
500- 59C
390°C
-420°C
Weight %
Promoter
'C
Weight %
Support
Material
S10-
f-
natural
S10,,
InfQsed
earth.
natural 1
molten
auartz.
silica
eel. Alc
o n«sn.
pfirrelaln
. T10-
fc
Dlsoerse
alHrlr
acid
Weight %
Catalyst
V,Or-ei.22
to &.ta&
Porosity
(
Surface
Area
Gas
Flow
Rate
O.5m/se
Contact
Time
SO,
Cone.
q.O-q.^
f
•10-60«
7 *
71
oa
Cone.
19.5*
201
Conversion
Efficiency
qq*
q?*
Decreased
7B.5»67X
qq*
Reference
'• '
CA-62-P-Q848<
TA-62-lP778r
— sTT^B —
CA-63-14366h
CA-64-1394c
CA-61I-P-E80S
Remarks
kcal/nole. Activation Bnerccv
23.6 Kcal/mole.
Bayer dual-catalyst orocess. •
Economic it tech aspects
V catalysts were determined.
OD t I.IDUI& condirt I.OZLS for
nvIHAMnn «ram> ri»>t«»nn1 nori .
CataXvst was oolsoned in the
presence of 0.33* HF and
rpRul tft Atuiils*ri.
was prepared. and tested^
.
o
(o
z
H
O
ZJ
m
o
i
o
o
TO
TJ
O
O
•z.
-------
CATALYST - VAHADIUH
^0
31
32
3U
35
36
Catalyst
YfuiflflUirc
Vanadium
Vanadium
-8-5-
V-0_
VATlAfM llffl
;vno_)
•*
V o~ (BAV
J .
Type of
Catalysis
Solid-
gas
H
"
"
n
"
Temp.
100-57C
UOO-62C
370-UOC
^ftn.Rnr
482 «C
Weight %
Promoter
•c
'C
'c
•r* K2°
Na.,0
K
(K/V=2.5/a.
Weight %
Support
Material
Silica
)
Weight *
CatalyH
V_0n=6 . 5jt
^
Poroiity
Surface
Area
Gas
Flow
Rate
—
o.fl-l. l
Contact
Time
SO,
Cone.
16.1*
7%
8.5-
O C4C
lOtf
fl-p?*
Cone.
16.7*
19.6*
11J&
IS-ifi*
Conversion
Efficiency
98*
max. 70j6
c5O* •
8o-
-------
CATALYST - VANADIUM
37
38
39
no'
42
ill
Cetairci
Vanadium
Vanadium
Vanadium
;V90C)
VanAriliiRi
[VoOJ
vcor •
Vanadium
V3°5
Trpeol
Catalysis
Solid-
agL a
"
n
n
it
"
Temp.
360-115C
300- COO
415-
591 "C
430°C-
511 "C
'22 °F
100-1420°
Weight %
Promoter
'C No
C Ha, K,
Ittj, V8
sul fates
0.4*
6.3-10.6*
K-0
opt- y-3*
11.0* Fe
0.5*
Weight %
Support
Material
Silica Ge
None
70*
silica
kipHplcuh
or celllt
silica
Slllea
SDherea
(3/lo"Dla
Silica
from wace
glass
Weight %
Catalyst
lb.H-33.9!
vco-
6-9*
6.7-7.7*
ope . o. f)b
)
••
Poroshv
Surface
Area
1QO-2OO
m^/R
Gas
Ftow
Rate
1 /Din.
linear
vel=2t
cm/sec.
Contact
Time
0.09-
O.U3sec
OPC.0.1<
sec.
SO,
Cone.
5*
8*
3.OV
3.27)1
°?
Cone.
20«
J-8S-
S.Bj!
Convergion
EHieiencv
»10*
81.7-99-6!
8U-08*
99-2-99-5
Reference
EI-52-1028
EI-S6-1O45
EI-58-1228
F.T-e
CA-57(P)-667(
Remark*
techanlsm of catalytic oxldn.
of SO*, with en^mfjrclal V
'article size = 10-14 mesh
V_0- alone is Door catalyse
fBr^SOo oxldn. but when mixed
with alkali metal salts good
results can be obtained.
Effect of Dhvalcal & chemical
factors in catalytic oxldn.
of. SfJp - Ineludjnfl pnlsnr^infl
3i catalyst, catalyst placing
ind converter design-
ni f fftrftnt eon^jnfl^ ti*on °^ ^r^r
catalysts were studied for J
the catalytic oxldn. of SO?
latalyst prep, described.
Sulk density. O.b2g/cc.
3t SO, wre studied in differ-
ential bed reactor-
Bulk density 0.7pe/cc.
sa-) with << kinds of feed gas
:onposltlon.
1 Preparation of catalyst
described.
O
i
o
8!
n
o
o
a
O
uo
-------
CATALYST - VANADIUM
15
16
Catalyst
V2°5
"
r0o_
gK^ . .. .
Type of
Catalysis
gas • .
ii
"
Temp.
7OO°F
Weight %
Promoter
Na,SO|,
iron oxld(
KoO-10*
p;o5-is&-
Weight %
Support
Material
s
Silica-
alumina
Weight %
Catalyst
V^-T-Sil
Porosity
Surface
Area
Gas
Flow
Rate
It, 320
ftVmln
Contact
Time
SO,
Cone.
12*
o,
Cone.
sat
Conversion
Efficiency
98-9956 •
Reference
- «
CA-'SU-aWSi
CA-66-P-1I550
_)
i
1
. . J
Remarks
catalysis.
High te'mo. corrosion bv V_0_
is decreased by the addn *" J
of dolomite and (NH|..)^SO,.
/
sbxidation of SO^ to SO-
'with low ignition catalyst
is shown.
%
i
'
I
\
o
z
(A
Z
o
m
55
ni
o
X
o
o
a
•o
o
a
O
Z
-------
CATALYST-VANADIUM
17
48
19
50
51
52
Catah/M
Vanadium
£eslite
V.O,
'3 >
V3°5
Vanadium
:mpds.
?pab le ol
elding
ranadyl
aulfate
Ikali
ictal van-
idate or
ranadl te ,
ranadyl
lalts.
anadal
lulfate
ir NH,, var
idyl oxalf
tan&rfj £
>nhyriT>1 rin
ietal van-
lOl. van-
ranadate
lolutlon.
Type of
Catalysis
Solid-
Gas
n
"
n
n
.
te
n
Temp.
100° C
itoo°c
150-500'
Weight %
Promoter
Salts of
in, Bl or
bi
Weight %
Support
Material
'0.11 1f»
Silicic
ic id
Stannic
tvAfnw^ rip
Eel
'umlce.
cleaelguhi
Hat.
:arth or
illlclc
icld
Iround
illlca.
illlceous
irlck,
land or
cleselguhi
md wet
illlca KB:
'Inelv
llvlded
Lndlfferei
md blnde;
Weight %
Catalyst
t
Porosity
Surface
Area
Gas
Flow
Rate
iOcc/Be<
Contact
Time
SO,
Cone.
'.2-T.01
°»
Cone.
Conversion
Efficiency
1T.85-39UJ
)«-95!l
Reference
•
JA-23-P-910— 1
:A-24-5585-t
:A-25-P-5b5-3
:A-25-P-131t*
:A-25-P-104W
:A-2S-P-2817"<
Remarks
catalyst poison te.tt. HC1)
passed over Pt-free V- zeolite
at elevated temperatures . "
Greater catalstlc action than
V- zeolite.
Preparation of catalyst de-
scribed.
Preparation of catalvat de-
>c>*lbed. Treated with ocirt
ras such aa SO in th= nni^
'reparation of catalvat do-
icrlbed pellets may be forned.
.,
Prjjinnrtit-.-lnn of aotalyet d»
crlDed.i^jTn^^^o,.^ +^^0* .
In the colrt -i^h • reducing g»
ntll V la conmletelv ndueed.
o
I
o
PI
w
5
5
o
z
o
o
a
TJ
o
a
H
O
-------
CATALYST-VANADIUM
51
51 '
55 •
56
57 .
Catalyst
Sn OP Bl
vanadate
Sri vanada
Fanadlum
catalyst
V-coraod.
Vanadium
Type Of
Catalysis
ii
e "
Solid
Gas
n
- "
Temp.
Actlvat
300-350
Oct. at
rate—
II90"C
Weight %
Promoter
C
Weight *
Support
Material
Pum1 !•*
as best a
silica Re
nl iinvt na
Eel or
HgSOa
noh-
pennuto-
aenetlc
contact
mass conui
active
silicate
Siliceous
diluent
and a
silica
hvdrotcel
11 ot-nml f.B
fragments
Weight %
Catalyst
B
Porosity
Surface
Area
Gas
Flow
Rate
"Soace
lelocitj
\ir 225-,
1.500 S0't
)-i!2
Contact
Time
SO,
Cone.
°7
Cone.
Conversion
Efficiency
-
»"58. •?-
QB.7*
Reference
»
CA_3t;_p_unQiV
CA-yfi-KKH-^
:A-27-P-110M
:A-27-P-225-
Remarks
Preoaratlon of catalyst de-
scribed.
Afc hloti 2O? patea .01* can low
tifi^P - obstructive* nrtnnT^ljloi^
nay occur. Insensitive to as
poisoning.
Elevatpd T^np
1
'orinlna metal
crlbed. Activity of catalvar.
ccsa and to oartlal .reduction
;o VaO(|J which acts aa an 0
artier.
o
w
z
H
O
m
(A
m
o
X
n
o
•a
u
o
O
Z
ON
-------
CATALYST-VANADIUM
59
61
S3'
63
61-
.-
Catalyst
Yflinflt'* um
V2°-
la Vono/lo
F3°5
from
CjVO,,
la Vanadai
Type of
Catalyss
•
n
e n
Solid-
! "
Temp.
iliruRno'
'
360-520'
158-198'
Weight %
Promoter
•
;
: K,SO,,
Weight %
Support
Material
nn
e
S10-
• *~
(1 A«Al gtlhl
uid gun
Silica Oe!
S10,
Weight %
Catalyst
Poroiity
Surface
Area
Gas
Flow
Rate
Contact
Time
SO,
Cone.
39V
°»
Cone.
f!lJ
Conversion
EHiciency
Reference
«
• ji_99-2fi68— 8
!A-29-7775-6
:A-io-F-i4Ssr
:A-30-P-31B25
JA- 30-3306-2
Remarks
n«1rlat*.1nn pBT> mfn nnff pgp ga
gf c&talyst la 5 tinea aa rapid
aa In a aaa contK. 7-81 SO.. •
19S 0, and 73S N,. Preparation
of catalvat deacrlbed.
procedure deacrlbed
Effect of varying porosity at
f*,l ftf A **««*• t:AMpa>**n*ii«>A* «*-fla»
scribed with Incrflflflln^ ftflrgpit}
the tenm . of ^hfi max. vleld
energies are the same for
<11 referent oopoalfclea hut: in—
Preparation. of catalyst de-
neplh>*ri-
-,
Preparation of catalyst de-
scribed.
CatAlVtlA AKt-1v1fy «lth And
without adding of S1O. ob-
served. Dissociation 'pressures
of sodium vanadate and oxidation
of SOp •"»« DMkAaiiiwrt at- -u ff np-
ent teams .
o
v>
z
o
m
o
o
a
TJ
o
a
o
-------
CATALYST-VANADIUM
05
««•
67
68-
69-
70-
7?-
CataVst
V3°5
from V
iontalnlni
ore
v n
ffoa
-TO,
Vanadium
catalyst
v,o_
•*
v,o. •
v,o.
*" •*
«• •/
V O
C 3
Type of
Catalysis
n
(i
n
n
n
n
Temp.
ll2«5°C
H50°C
'Weight %
Promoter
In 2
Cone.
17*
Conversion
Efficiency
Reference
1 *
••-30 p_57£n^
:A--»p-p-infi»
:i_iu_in?R-fi
:A-T 4-61^6-6
!A-3ll-71fig-ll
:A-114-8l8R-S
CA 39— 910 9
Remarks
Treatment of V-contalnliuc
material, to render It
catalvtlcallv active, de-
scribed.
material described. "
by Fe_JSO,,)., Investigated.
Mechanism of oxidation of SO,
postulated.
E J
Acfclulf-y f\f v n fVunri t*o bo
leas than f-.hat xp run NA^O. T
V->°c Hltfl SO-, -to Dive unnarivl
BulfAt* (9;iS.n__'_: „„ f J
* »»2gl«en.
_ >•
o
v>
e
ro
o
o
a
•o
o
H
O
z
-------
CATALYST-VANADIUM
73
TA-
TS
76
77
Catalyst
vr°-
^ •*
AJC p«ta-
rina^ate .
3°5'
r~o; + Fe
:r-J0_
• e-*3 —
v,o5
VgO,. §Pt
ir^V-O- +
Pt £ '
Vanadium
2 ™
Type of
Catalysis
Solid-
iflfl
n
n
n
ii
Temp.
420-480'
tlO°C
41l|0C
I|05°C
120°C
438°C
39 2° C
188°C
123°C
«20°C
110°C
Weight %
Promoter
KflH
Pe, Ca
3nO,
: Pe.
!b« 0,
K20
Bad
Weight %
Support
Meteriel
Illlca KB:
md infua-
irlal earl
n
HO, with
int fYpnu^
illllmanlt
>onded by
illl cats
tubes
Weight %
Catalyst
i
!.
Porosity
Surface
Area
Gas
Flow
Rate
Lin . vi
0.15m/ 8
^7 . ^ em
sec.
Contact
Time
L .
c
' Ol
SO,
Cone.
Opt 7-
»
Opt 7-
B*
T*
r>°*
75*
90S
951
50*.
251
11
8. 51
U}J<
°»
Cone.
19.6X
10. 5*
25>
101
51
50%
751
111
12.51
181
Conversion
Efficiency
Reference
•
•A-UP-KHQfie
:A-42-8058h
:A-'(3-2711e
: A- A 7- 80 7 He
lA-llT-BP^ls
Remarks
catalyst and method of
activation described.
Oxidation studied with differ-
ent catalysts, terns, soace
velocities . and comDosltlona
of gases. The AK vanadate
and the V O + Pp anrf Pn
were beat. J
vestlgated.
fethod of preoarlnK catalvat
lescrlbed.
^DDfira1:ur@n yl v^n ax*e «-.h«
Lowest tenmaf^ture of *h« g»«
Phe extinction temn. on rnnl-
LnK Is lower than the honHnublllflr •ftnrt'T
O
v>
z
O
m ,_,
w I-1
pi a\
> vo
O
O
a
T)
O
O
-------
CATALYST-VANADIUM
79
80-
81
82
83
84 •
85
86
Catalyst
v n
tf i?
IVOj
Vanadium
V.,0,
*~ J
Commercla
V o
c J
Vanadium
catalyst
Vanadium
catalyst
V,0_
Type of
Catalysis
n
n
Solld-
Qas
it
n
n
it
Temp.
•Ann-
1000°P
• HHO-
l|70°C
360-
licnop
•508-
536"C
450°C
Weight %
Promoter
•12-22*
Kn°
^
21 Ha
BaO
BaO
• T103
Weight %
Support
Material
leg^lf^uhr
•10-55*
310-
Peliets
Silica Ge
Al,0,
Weight %
Catalyst
•7-12*
V.O,
12* V,05
"
Up to 25*
at 81
V2°5)
Porosity
•60*
Surface
Area
•4.8-
258 cm2/
E
Gas
Flow
Rate
•0.65-
sec .
•2500
1/1-hr
1)00-700
Contact
Time
SO,
Cone.
•6*
7«
•6*
L7-80*
•4-7.1*
oa
Cone.
•15*
4-18*
Conversion
Efficiency
•82-92. "5*
••vSO*
•37-78.8*
_
Reference
•
ra_lifi_P-fiR7fi|
CA-46-P-9902I
CA-47-F-67l4i
CA-48-F-9S86c
;A-48-8003i
3A-32-3687-1
CA-48-13386f
na_liQ_in7i7(f
Remarks
Fluldlzed oowdered catalytic
mater^{\^ Ascribed.
Aooaratus uslrut fluldlzed
vibrations described. 'Drain
81 Zfi 0.^— Qg^nm
Effect of vibration on
flow described.
Preparation of catalvat de-
scribed.
Reaction rates det,ffnn1 nf*rf anrt
mechanism iri v**r^
adsorption of HP i^veatiTi,f.Prt
Regenerated calayst «houA/4
Increased activities.
^ate constant dependence- on
5 articles demons t?«a ten _
-xamlned at vAiHnna v n ' ^rtn
o
2
H
O
TO
m
M
ra
>
n
o
z
o
o
a
•o
o
O
-------
Catalyst-Vanadium
AT
88'
89'
92'
• 93
. .
Catalyst
V [.
2 5
V O
E '
V-0-
catalyst
v~°5
Connercli
^^
Type of
Catalysis
n
n
n
iolld-
L "
Temp.
«n Kin
250«C
360° C
380°C
»25°C
»55°C
>50°C
'380-
150" C
114 5-60°
•37K-
Weight %
Pi omulsi
P
HIBl^O-j
n
Without B:
H
n
•?f?o/v 0
• 4.S/1 to'
i.4/1
la.,0 10. o:
[ffO 3 . 19
'
•K8SO^ 33j
Weight %
Support
Material
Be water
Klaaa(17!
SlOo)
•QuartE
son
310, 58.11
•Silica
Weight %
Catalyst
8* V.jO,
n
1/2 05
Porosity
Surface
Area
Gas
Flow
Rate
oer 1
•Bt-.nlyRt
iDace
>er hr.
•mass v
Lb>hr-BQ
Contact
Time
3.25 ae<
n
H
n
n
n
1.2-9.2
1. «
.ft.
SO,
Cone.
7%
n
n
•
it
n
•6.5-
9.61
7*
•0.2-8.(
Cone.
t
Conversion
Efficiency
90.21
98.8s
25-30!
66.01
98. 6S
981
Reference
' *•
*&— ^0— ^0"5Qp
*A— 50— filfi^lb
IA— 50—P"l005?
*A^51"12ft^5i
!A-51-17023a
Remarks
WhAn V ft tn vwriiif.»H hy Qn
at low'tSmp VnSOn 1n f1 >>af
fDI-mOrt <-hnn»1no- ?n V ft if
higher temps. V-0_ la^f^rned
directly. ~ " •
Hydrodynamics of gas flow In
fluldlzed catalyst system
calculated.
curve of late K v^ 1/T i,g *^»-
catalyst to the Inactive V
aulfate.
mixture la descrtbBrf •»»!•».
Ine K^O : V^Op rattn rnv* Iminr
Inltlatlnif't-ompo
9fl^aAyat dennrlhpd.
Coraparalon betiiean f i »«rt .^
Fluid bed 1« lie f-J~. non.
effective, (•eafcaivnt £r^-8o
mesh).
O
z
u
z
H
O
o
X
o
o
a
•o
o
H
O
Z
-------
CATALYST-VANADIUM
oft
Qfi-
98-
qq.
Cetelyst
e j
v n
^ 5
v,o5
v2or
*2%
_
Type of
Catalysis
Solld-
n
11
n
Temp.
ipt.-les
han 44(
3on_Uanc
180-5501
ft 50-525
Weight %
Promoter
'C
'. K ^n
i H
Weight %
Support
Material
Ueaelgulu
md SK)
:,0'5-5-8)8 310,
£-6
1. 55-1.75?
K,0
MA H
£
: zNa.,cr
"
S2.S-62.^
Weight %
Catalyst
^,0
L3-5-7.5J
•
Porosity
Surface
Area
Gas
Flow
Rate
Contact
Time
SO,
Cone.
7*
°»
Cone.
Conversion
Efficiency
Reference
:A-5?-lliqQlla
•»_•;?- i7fi^7b
IA-S1-P-11781
CA-53-13755B
Remarks
influence of pore-size and
on activity of catalyst Is
discussed.
discussed.
PhfttG CQntfiltAff71 *>f active
V O -K SO^ mid V O -K S n »
A€ tfernn between Sqrf-U6nsn'
described.
CffDlDPT^rt- rontaot: MmA sn
ntudlari
described.
2-8 ten reaction fcubea unori
The rates ofLrea'ctlon at
Pld diRpiina^d.
.^CtlvltleR g1 von
-
O
u
z
O
a
O
I
O
O
a
•o
o
x
5
-------
CATALYST-VANADIUM
ini
10!
10)
10'
10*
lot
CMSlya
VnnnriliiHi
Vanadium
Catalyst
K
Catalyst
VnnaHl nm
catalyst
•(BAV am
SVD)
Yvi&ft* uni
catalyst
•_..-- - - .-_• .
Type of
Catdrsis
H
n
n
n
li
- .
Tamp.
t?sn—
380-550'
185°C
140-520'
;oo°n
450-490'
Weight %
Promoter
'eO'-17.«)
u.c.-o)
;
•
Weight %
Support
Material
&1*O»
gel
U-0-3'2tJ
Slug
>0-80J(
i!03 or
lr-y H nrtTH ,
Iranules
Weight %
Catalyst
IV (as V O\
7. 2-13. 5V
»1
Porosity
•59— 75fl
Surface
Area
•32=A7_
Gas
Ftow
Rate
^2QMU
1.4-0.9
i/aec.
!Vol 15(
'npl-.lmun
.700
I 1-2 5
i/sec.
0 S-1 '
,/mln.
Contact
Time
10.07-
1.15 sec
)
SO,
Cone.
•0.5-
7.0»
2-181
7-30%
71
2 7-7 '(
)
°?
Cone.
'10-701
Conversion
Efficiency
'up to 931
101
13.7*
'80.6-
98.1*
>80-40X
LO-70I
Reference
•
;A-53-20624g
:A-5*I-^877c
!A-^ll-P-
IA-54-219761
lA-SS-^Olle
lA-^S- 11^26
Remarks
A1-0- Increases the thermal
by stabilizing the ohvslcal
structure.
•Catalyst Qraln alze l-2nm.
Calculation of velocity con-
stants of oxidation described.
M^fchod nf pIwpAlHnir ophA«*4j*^l
concentration An^ vA^iAt-jnn ^
Of yield With hel<7hf nf
diameter, and voluim vi"ln"1*'j'
Is discussed
•Particle sl'ze o s— i <*mm
Effect of particle fllza on
ictlvlty In fluid! zed bed
llSCUSSed'. Rate nnn.t-.x.
ipproz. 30/aeo for nart]f«.i^Q
•apldly for nart1clo« >i cj-n
n dlam.
•atalyst narf-i^ia a«.n n 70
:ratlon and tenp '
I
a
8!
m
>
o
o
X
o
(JO
-------
CATALYST-VANADIUM
i
o
I
o
o
J°
101
10!
lit
111
112
. - -. -_.
Catalyst
v,o_
Vanadium
V3°5
Yftn&dium
catalyst
"BAV"
v,o5
..-..---.-_-_. -•
Type of
Catalysis
Oas
n
n
11
n
n
.i.
Temp.
ll85°C
ISO0
Weight %
Promoter
Sb^S fl
CS-S6V
C8?sJ,
Rb^SO,,
CB -SO ,
).l mole
J..SO.
).? mole
la,SO,.
1.1 mole
i3soh
1.1 mole
16,30,,
1.1 mole
Cs_S20_
Flth and
ilthout
la-SnO-
2 3
Weight %
Support
Material
Siflj
Kleselgutu
LI 0 -S10.
E 3 •
Weight %
Catalyst
3-101
L molp
f-0.
"• ^
Porosity
Surface
Area
Gas
Flow
Rate
Contact
Time
SO,
Cone.
7-11*
o»
Cone.
Conversion
Efficiency
17*
Reference
- m
:A-SS-2^000h
:A-S-5-21011c
!A-Sfi-gl|fi8c
CA-56-ll972f
CA-57-3247b
Remarks
described.
Method of fyflnn pmppt-em
and catalysts described.
Promoters mentioned MVP o.
times the catalytic activity
Klven bv K sn.
Catalytic oxidation nturfioii
PrOlDOtlnR action nnmpni*^ fe\f
different Allrnll HIA«-.»I nnl_
fates using 0 Isotooes.
4 mefchrtrin n^ Aa^aina*. Dwna— •
LnK the A^vsr\tfifffi °^ fch*
flu-lrflfAH hAri ridRptHhod
.
-------
CATALYST-VANADIUM .
11 -
11 «
11T
118
, 119
Catalyst
Ymflfllmn
catalyst
Stgfig
v n
V3°5
"BAV"
V 0
£ 3
f,o5
V»9
T»peof
Catalyst
Gas
•
n
H
"
"
Temp.
188-
45'C
iOO°C
QO-4100
•Wear
Resis-
tant
Ignltlo
T^rw-17
•BAV-
ftio°c
I-JO-MO:
Weight %
Promoter
jijSO,
HfijJl.SO,,
K-S-O-
SOIST^
wn
3
•
x r> •> Bt
°c
K.O 5.1*
""
Na,0
c
Weight %
Support
Materiel
310- •
rTVhr
»1,0,-S10
Weight %
Catalyst
IS.8S
v n fi i ^
^ 5
V_0~ 6. IX
PorosHy
10ft
Inner
aurf .
120 n/K
ope vol
^fi5££^
finn_Rnn
•BAV-
i^nn.
1500A* '
1100-
1SOOA0
Surface
Area
Gas
Flow
Rate
0.17-
1.1 m/
sec.
fint/npp •
4500/hp
•n 31-
0 . 19m/8
i ?f>f)/rir
Dllot
iliine
Contact
Time
c
SO,
Cone. -
7-30X
.3»
•T ^-91
°»
Cone.
•11.5-
11.8j(
Conversion
Efficiency
SM 1
ton
imm17ig7r.
Remark!
described.
Industrial scale fluid! zed
bed described.
Catalyst particle alses of
0.127- 4. Son compared.
efffcts were lnvpfti-.1ffat.jid
UBlrm small glnnn Trnr.r.nr
fluldlzed bed. , •Drain B!»
l.^nm dla.
described. Vgr>]()un mlltlirrn
cnmflAro/1
•Effegt n* p n- J" 11 "np Tm
VjCL Has the atartlnic material.
'
I
3
m
|
o
o
a
a
-i
5
-------
catalyst-vanadium
121
123
124
12^
12 «
125
12fr
Catalyst
KVO-
3
v.o.
Vanadium
catalvst
va°-
V3°n
'S"y
v,o.
"SVD" com
V_fflt-a1yA
Type of
Catalysis
Jo lid-
las
n
n
n
n
Tamp.
"100° C
185°C
IOO°C
?80-1SO°
140-110°
"460-520
Weight %
Promoter
K.SO.
K,0
: K^SO,,
~
c
Weight %
Support
Materiel
U-O-j-SiO,
tieaelguhi
Calcined
lydroicel
110,
"
Kleaelguh
Weight %
Catalyst
'30j 7.2*
loles V,0r
4*>T* IDAl*
U0« 1/3
.-
.•
Porosity
Surface
Area
Gas
Flow
Rate
^Ooi'/hr
nnncp
Tel •
1000-
8000
nn_9nnn
|j|^ ^i]i1 n r
Contact
Time
SO,
Cone.
IDS
0.2«
•-2.S*
7_i nf
°»
Cone.
20(
2S
1 1<
Conversion
Efficiency
ITS
)^X
[T»OlIp"lH
Reference
»
!A-fiS-Q8mid
:A-6?-p-
-11835f
!A-65-P-
-19352d j
bA-66-79958n
70.5*
ung round
20.5*
1591
IS1
IA-66-87175V
CA.68-P-
-l42QS4D
PA-ftft— il^^^fif
Remarks
Sulfur! c acid produced at max
Load capacity of 10 fee tU SO,,/
Sg catalvat/ day.
Preoaratlon of catalvat with
addition of natural or
synthetic flbfir tQ ^ncrease
oorosltv and realstance to
fibrw*^ nTl 1s d^scplbpd.
S1O. carrier icround In a
vlbrnfclnn mill.
SO, precipitated from flue
fcaa aa (NHi.)_SQL.
SO. recovered from flue fia.9
as (NH.. )- SOt fool nim>).
cuaslon of the reaction are.
presented and V catalvat com-
PRre.0* with Ft eat^yst-.
01
z
o
21
0
m
O
X
o
i
3
a
-------
CATALYST-VANADIUM
12
12'
130
13i
*"
HlJ
catalyst
V-Catalvi
• f?Snf» •
»a°-vA'
V-catalvf
Ca-V and
Pe-Sb-V-
c&talvstE
iggj
V 0
V.rntalya
_
Typed
Catalysis
t Solid-
Gas
t n
n
n
H
It
Temp.
>UI40°
. .
Weight *
Promoter
•
7»m 1 f»
Weight %
Support
Material
J^|f*f|««lCTl^h
Weight %
Catalyst
^
Porosity
Surface
Area
Gaa
Flow
Rate
Contact
Time
SO,
Cone.
°»
Cone.
Conversion
Efficiency
Reference
• *
!i_p8-m«i»
•
IA-28-P-2H787
:A-37-61l67
CA-*»2-39l8d
CA-22(P)-
1018*
Remeriu
The Influence of contact tlioe .
PC^TDflfll t*on °f van ml rtllTfl
nnd temoerat111^ on the
Reflultfli (•Qv*iM1*'kd to thioBe
w1 th Pf-
Prooertles of rf*^f|^yitf-H
studied.
scribed. 3
coefficients for BeT^ml
nxldntl on of* Jtfl|
*
1 n fl iiAn<*A nn n*\«**\tt« <*B«*n4 •*•
catftlyatB.
Method of PPCD^*1"^**** eontngt
-
Ul
O
Si
5
a
o
z
o
o
X
•o
o
X
•H
5
-------
CATALYST-VANADIUN
IT
1^
137
.139
140
Catalyst
Vanada.te
p** vann-
dltea or
?e, Cu,
Aa. "l.
Co, Tl.
Zr . Ce .
Al fin
CA U TT*
or Mn
Zeolite
V.
V-compds.
V-catalva
»A hv Pr_
eatalvst)
b-CatalyaJ
Type of
Catalysis
Solld-
— J
n
n
n
Temp.
•410-
•450-
fiflfloc—
•700-
Weight %
Promoter
•Alkali
Zeolite
InR K
Weight %
Support
Material
Zeolite
ing. Al.
Re, Cd,
Zr, Zn or
Tl
•(33° Be'
Cellte.
Kleseleuh
etc.
Zeolite
Ing S10,
and- cata-
IvtloallT
active
metals .
Xleaelguh
nnd SID—
[laucoell
•V-pellet
Pt-Abesta
'
Weight %
Catalyst
•5-10*
•Zeolite
V + other
catalytl-
cally act
Ive metal
9
'
••
Porosity
.
Surface
Area.
Gas
Flow
Rete
Contact
Time
SO,
Cone.
• 7*
o,
Cone.
Conversion
Efficiency
Reference
•
486^
CA-23-fP)-
940Z
CA-25(P)-
10H57
3O336
CA-27CP)-
3A-^1_^3179
Remarks
are olven.
scribed.
described.
-
Pt catalysts.
Wlfch frhnt nf P*
o
o
•s
> -^
s °°
n
S
|
>
6
-------
CATALYST-VANADIUM
lii.
HV
14V
lllfi
117
ins"1
Catalyst
ITgOj
V-catalva
Y.O,.
' °«
2 "
'E°-
Ba-Al-
Vanadate
V,0,
•"
Type of
Catalyse
Solld-
gas
n
a
n
n
n
Temp.
400-
550 °C
1&5-550'
!i3i
Weight %
Promoter
Ha 3OU
f "*
CuSO^
»o rsn ^
BaCl.
K3S01
BaO
»i n
Weight %
Support
Material
MnO... in-
fii^pplal
:laln. a-
ibeatos .
lias 8 ,
luartz
land
310.,
Weight %
Catalyst
Porosity
Surface
Area
Gas
Flow
Rata
Contact
Time
SO,
Cone.
k- ill |
°»
Cone.
llfiS 1
Conversion
Efficiency
it> to 97. 1
Reference
•
Ck-12-19613
:A-l)0-29ll38
3A-57-1101
CA-29-26995
JA-35-33967
!A-^'a.T7U7h
Remarks
In aulfurlc acid production"
discussed.
oxide catalyst discussed.
Klnetlcn 4* ""yflsed. '
scribed.
Effect of promoters and
carriers on catalytic ef-
lEffect of v:i^1oua Ann7ujltr1 °^
iK,SOj, and sio. on ca^ajiytlf
activity of V O In rtl«-.i«=.rf
i ~y
lEffect of S conec»nt.i>At-iQp ]ri
studied.
Renction valnrltlnn ralnuleted.
u»
z
O
I
o
o
a
T)
o
a
-------
CATALYST-VANADIUM
150
1511
153
153"
154'
155'
157
158'
159
ifin
Catalyst
/-.O-
"• •*
V-0_
Ba-Al-V
V,0V •
r-catalys
r,0,-
fc •*
igfij
1AV
K-Vanadat
BAV
Type of
Catalysis
So 11 ti-
ii
n
11
n
»
n
n
» it
n
Temp.
,cnnon.
420-525
>5QO°C>
170-608
1)00-500'
5 WC
?«; «;r;no,
Weight %
Promoter
08,50,, '
•4-6X
BaO .
Al O
" 3
C
c
Al.O,
~ ^
BaO-Al.,0..
2 3
C,SO,
-
Weight «
Support
Material
Silica Oe
Silica g'e:
illlcon m
Weight %
Catalyst
61
1
Porosity
Surface
Area
•<0.05
mz/s
J.5-21.!
n!/a
Gas
Flow
Rate
Contact
Time
•0.435
sec. '
SO,
Cone.
•101
75*
o,
Cone.
•181
25*
Conversion
Efficiency
78.5! '
18 ll«
Reference
'•
CA--54-ll672d
CA-54-ll672e
CA-54-2196lf
:A-55-15139e
:A-56-2025h
:A-58-960d
I A— 59— ^?48b
•A-59-458le
:A-fiO-2f68f
IA-60-3527C
Remarks
Results of thermofT-anhle
study presented.
Effect of alkali-metal sul-
fates on catalytic activity
studied. 'Velocity constant
9 times ffreatex* than brl r.h
X,SO,,«klnetlcs discussed.
• "1
Particle size 1.5-2. Omn.
discussed.
Mechanism discussed.
Rise of i-eslatlvir.y nf
catalyst due to A1C0- ex-
plained. Mechanism 6r SOe
oxidation derived. • "~
•Poisoning by As.O. discussed
nt!i,ri1»H • • ^
I
I
a
m
w
m
n
z
o
O
a
•o
o
(DO
-------
CATALYST-VANADIUM
16?
164
166
167
168*
Catalyst
darn
SVD
V,0.
V,0.
f-catalvs
V£0r
r£°5
"
^
Type of
Catalyiii
Qflfl
-700°C
n
ii
n
n
"
Temp.
485°C
4.8500
500°C
370-400
Weight %
Promoter
Kj-SO,,
11.5)1 K.,0
C
Weight %
Support
Material
Silica
fllMm^ nfl—
silicate
VKhKC?)
Quartz,
fused
a uartz .
marchalit
and In—
f i^aejrlal
earth
Al,0.,
K.,0 J
*•
sio3
part, slz
0.5- 1mm
Weight %
Catalyct
S.P-B M
,
8*
t ••,
Porosity
Surface
Area
Gas
Flow
Rate
2000hfl
Contact
Time
SO,
Cone.
o,
Cone.
18. QS
Conversion
Efficiency
Reference
*
n»nfinniTn?h
CA-6l-10082e
CA-fil-P-
- 11376c
:A-63-7S25d
7A-63-P-
17527h
*A— All— P— 1 M3h
Remarks
scribed.
carbon black examined. At
high terac. C700°C) C cauaea
when }(B o^t >51-
Data correlated »lth the
average oarticle alz*> and
the catalyst auoDorta .
New catalvnt: rt»«o«lK«^
described.
EffectivenesB of vanimiri
catalvflfc Rhof.on ^«»«7rtc5r*d
D^Rlffn oT nni,4nn«.^ *• SO_
l*etBOVnl Ttvim T1 no p*Pl rtrtfl*"*1 11—
ed.
i
H
O
X
PI
I
O
O
X
H
O
z
Co
-------
CATALYST-VANADIUM
170
171"
172
•1711
175
176
177
178
Cnalyst
V-catalys
V-catalys
V^O,.
~ •*
V Q
POV
(Ba-Sn-V
V.O-
(ZK'O-BaO
•0.5 Al.,0
•V-0--12
310,1
V-0,
f.,0.
Type of
Catalysis
11
N
n
n
n
n
n
Temp.
480-550
440°C
liso-C
>440°C
470°C
420-554
C
240-383
C
Weight %
Promoter
KVOv+K.,SQ,
K3C03
T1O,
SnOj
BaO
K n
*
K-SO,.
10* K,0
Ml L
Weight «
Support
Material
alumlnn-
Silicates
K-sillcat
silica
gel
Kleaelfcuh
KleselRuh
Weight %
Catalyst
v n
' *
8.5» V,0
Porosity
Surface
Area
Gas
Flow
Rate
Contact
Time
SO,
Cone.
4-23S
k-6jt
°2
Cone.
6.5-UOt
t-20*
Conversion
Efficiency
21-91%
<22I
Reference
'-
:A-64-769Cfc
CA-68-63031 k
CA-68-P-
5310lh
CA-62-3445f
CA-PO-6171'
Eft-32-10551
:A-51-6lOOb .
:A-61-B942a
Remarks
different sources Stddled.
Design of equipment described.
Structure of wear resistant
catalyst studies.
that have hiKher activity
and lonirer life.
Catalytic activitv accelerated
Use of Barium Tin Vanadate
Catalyst described.
.
Iraln alee o.67_i sfim-,
ihOW that the Jlnpopppf-. r*al-«i_
ihemlsorbed SO. >nrf ^v,.miBnT*ort
Itomlc 0 to form <;hflpinophi.d
SO. Anri ft VAcnnt. nltl*.
I
O
X
8!
O
O
a
T>
o
z
>
5
-------
CATALYST-VANADIUM
Iftfl
*181
1*i82'
181
4KL85
Catalyst
C '
t o
'e 0
>t? 3
' 0
f *
I O_
" J
UK vana,-
date
'z°3
3CB vnnn-
11n eata-
IflTSt
Type of
Catalysis
Q&8
n
it
it
it
n
Temp.
C.
450-
500°
440-580
325-
550°C
6^0-800
Weight %
Promoter
"-!« "a
Na 0.15*
R 101 Na
2.9%
K-»0
ta.O
also Ctu
:r, Al,
(. U. Ca
Ions
C
10.6* K-0
Aa
P
Support
Material
iio^
KleselKUh
we^rti TQ
Catalyst
^ 3
tf.O_ 6.81
7.7%
••
Surface
Area
Gas
Flow
Rate
.2-1/hr
inono
»*/hr
150-700
1300-
3400 cc
gin 1.5
3.1fos
Time
Total
(5 Btagi
1 On nor
0.008-
0.039m
SOj
Cone.
8*
;-7*
iT
S-12*
1 3*
1.
"j
Cone. •
5.2-13*
Conversion
Efficiency
66-98*
071
97*
57-97*
58-61*
Reference
'-
•
'A-24-30SQ2
EI-61-16Q9
:A-50-9211
:A-23-1725
3A-65-2892D
SO oxidation or| fl^^H-
&f*4-4u1t-v nP vnv-4n»? rniTffH -
nations of catalysts and
pi*omat:*iT**i Inwnfc'f ^ffltC^r
nr»r>Hur*H nn oT H Sfl^ *^^1Ca7ib~
6d. enDlOSltlR 5trntnlynt
lAvPvn ranaeltv H SO./
opt*. 3 AhT*a*i 7 fcnna/t.sin
?fev^.ew of kinetics, mechanism
ftnrt f.vn*»«i nT />nnv»f *•*>*•* "
Data fflven fos* ^^p^^'TipntS
Activators In the reaction
uslntz V 0 p 00
s w
n
o
TJ
o
5
-------
CATALYST-VANADIUM
*1B6
*187'
1RB
*1BQ'
"*190
4fl91
"^102
"*19V
Catalyst
.nduatrla
anadlum
catalyst
Vflnfid.iun
atalyat
v n
i 5
satalyst
ru°-
J°5
V 0
PStasslui
anadl te )
Yflflflfl1 V"l
catalyst
Type of
Catalyse)
soiia-
u
n
it
n
n
n
II
Temp.
(IM2-OUU
180-^60'
(SfiOpp^
120-
550°C
uoo-sso'
TWC
100-
iOO'C
iOO°C
Weight %
Promoter
i'
c
Sxchafif^f —
ible baaei
c
K 0 16 51
30, 27.3*
KO.*-?.*:
Weight %
Support
Material
!10- 39 6
fc
lvdrof*(
1050
Contact
Time
SO,
Cone.
8*
8*
7-1%
t-lOt
1.15*
IS
i.5S
°?
Cone.
151
Conversion
Efficiency
q7f
qfix
99.11
98*
)8.5*
Reference
m
:A-55-HQ»Jt>
:A-aa.«i7i
:A-2?-P-lQQ8J
:A-68-P-HS21
:A-51-13509f
louraal
Article f2Jl
:A-28-P-8678
JA-16-P-
-H608d
Remarks
iDtlmum oroceaa conditions ror
a multlstaffe reactor discussed
3.9? ton's catalyst required
broposed contfict-proceea con-
verter dealicn'
Preparation of catalyst olz-
tureq described.
117-Wshort ton HpSOii/iay
r^eeesfla|^ In described 2.— 8tS/f*^
process .
Bench scale Investigations of
SOp removal deaei-lbed.
*•
Klnetlce and loechanls&'of,
reaction catalvzedt by vanA-
11 urn oxides dissolved In
Investigated.
nlxtures are described.
.
•Description or ppnat»9*n« p^r
Contact pfOCfnm nf U ^^t. P***—
ludtlon descrlb»ut '
>.
i
H
O
X
O
O
O
TO
O
-------
CATALYST-VANADIUM
*h .,
5 S.
a ^
o
i
o
o
•a
TJ
o
TO
o
z
-------
CATALYST - IRON
T
2
It
«s
6
8
Catalyst
Steel
Fe~0,
1 DVJ
oxide
Pei°li
•j'
•4
w
Gas
Flow
Rate
nin
n/aec
VUif*>*lUUI
D 07*?
a/8 ec
Contact
Time
i hP
hra
0.27 s<
•
so,
Cone.
n i f.f
12*
1?*
»12*
°?
Cone.
18.5-
1Q*
Conversion
Efficiency
7_in«
?6-l*0*
Reference
«
CA-13-5910K
CA-ln-8l25d
CA-SS-Qesb
CA-6l-9l85d .
CA-6S-l6'5Ii^d
EI-66-1OO2
Remarks
teatliuc ateel with Al conelder-
ibly reduces the rate of cata-
lytic oxldn. Ana material.
laxlmuin catalytic aet-ivltv
La between 580-635'C
Ln the fluldlzed bed at 6OO-
7oo"n.
:atalvtlc activities of ffleta^
>xldes for the oxldn. of CO.
:he aDDllcation of icethod for
the oxidation of SO?.
Itilizatlon of iron catalyst in
Kh» rnnf-.jirt. nhjkmho^ nfnfoan ^
>xidn. of SO- with a fiuiriizeri
>ed of Fe oxide ^n contact
imr^na.
of S03, may be used for contact
tower process in manufacrore
of aulfuric acid. And
material.
problems of usln^- frnn fBt^^yst
1n ^nnt'.jio f*. f*.mMi^ niw*0anA
Eauafclona are derived foT
kinetics of oxldn. of SOy .
And mn ^-n rlal .
-
o
z
CA
>
Z
H
O
TH
m
w
PI
n
x
n
o
7>
TJ
O
3)
O
Z
00
ON
-------
CATALYST-IRON
9
10
11 '
•-
12
IT
13
1
14
• 15
'"r
Catalyst
Pel OH) 3
Fe.,0,
Fe~0'-Cu(
PeiOV-
FejOj
Fe
Pe-iOo '
Pe203
ffc 0
£• 3
Type of
Catalyst
Solld-
Oas
Solid-
Gas
Solld-
GBB '
SoUd-
Solld-
Oas
§ol Id-
as.
Solid-
Snl1ri_
Dan
temp.
125-131
350-60(
640-80'
..
250°C
650-72C
Weight «
Promoter
'C
'C
c
Weight %
Support
Material
Pumice
asbestos
BaSOii
paper or
-
ter Gle
Weight %
Catalyst
Pure Pe_(
1-70. OJ~(
IS
Porosity
lo
Surface
Area
•• •-
720 cm2
/en*
Gas
Flow
Rate
31/hr
100-20C
1VO16 gf
Vol cat
- ~
S.-176-
n/aax
'?•
* i;7^O-
17190 h
Contact
Time
i/unlt
,-hr)
fe.4?-
0.586
sec
.-1
SO,
Cone.
101
7*
7»
10-14%
*7-12»
o,
Cone.
*9-18.9
"9-1*1
Conversion
Efficiency
in-71 lit
40-73.4*
S"i-lQt
Reference
-
p«_3l_(;R^n9
rA_?7-?iasl
CA-37-21451
IA-40-5211S
;A-44-P-27l6c
:A-6i-ii6in«.
-
CA-62-8fis8ff
..
:A-63-P-
CA-66-69286c
Remarks.
ThQlco oT nuppnr-t motflrlf1 «lth
respect to nan velocity dls- •
cussed.
Variation of ar.Hatty JQf nata-
lyst with variation of flow rate
temp, and I CuO measured at low
. .
Rate of SO-; formation and enercv
of .activation given.
• — •
PepOo used as finely divided
powdBr suspended in the gaseous
under Drei|Ri|T*^'
• ' ' "'"•>-•
Preliminary 'oxidation 91) flul-
descrlbed. Particle- slze-0. 80 -
1/30 mm.
Dimensional ecoqomlcp fti"^n-
. : . -
.the oreae'nce -of HpO discussed.
"'..'"
Fe^Oj and SOj recovered at 500°c
described PnT*r1il*»- d1«ni — O.7.^ ipm
o
z
(A
Z.
o
TO
m
m
5
a
o
z
o
o
70
TJ
O
O
Z
-------
CATALYST-IRON
17
18 .
19
20
21
22
23
24
25
26
Catalyat
Fe-,O,
Pe?°3
Pe,03
Fe20-,
Pe =
Pyrlte
p lnrf*»T*j»
Pe^Oi
Fe203
Fe50}
_
Pe,0,
FeoOj
Fyrite
cinders
Type of
Catalysis
Snl Iri-
Gas
Solid-
eas
Solid-
Gas
Solld-
cas
Solid-
Gas
Solid-
Gas
Solld-
Oaa
Solid-
Gas
Solid- .
Gas
Snliri-
Gas
Temp.
mo-uso
100-600
-
Rnn-fisn
Weight %
PrOmwIer
c
c
Cr.Mn.V,
or Zn oxl
c
Weight %
Support
Mswcial
Ag
es
Weight %
Catalyst
Porosity
Surface
Area
-
Gas
Flow
Rate
Contact
Time
SO,
Cone.
3-4*
7*
121
Oj
Cone.
.
Conversion
Efficiency
95X
up to MOX
93.44*
90.34*
Reference
•
JA-66-P-
87178y
CA-66-
-lOBTOOe
.
CA-21-25563
CA-31-4882T .
CA-6l-2522d
CA-65-9786d
CA-68-53714W
CA-68-t)OB02e
CA-51-9122h
CA-S3-3621B
Remarks
Process of SO? recovery from
flue gas described
inrj.uenee or Ag carrier and
doping with Pd and Hg on energy
of activation discussed
Method of preparing catalyst
given In supplementary
reference
Mechanism of reaction via
Pe?(SOii)ii ttlven
Stationary and fluldlzed Fe
catalyst layers rate constants
determined. ^Equations elven for
making various calculations
Degradation or catalyst during
operation discussed
Radical .theory of active oxides
discussed
catalytic surface discussed
Oxidation of SO? catalyzed by
Fe2u2* Cr2°3 ana. coal ash, fire
clay and glossy slags In combus-
tion space of slagging bailers
investigated
3y products of lUSOi, prod.
Jsed as catalystT
3.
O
z
Ul
z
-I
O
TO
5 M
m oo
> 00
•a
o
o
o
a
u
o
a
O
Z
-------
CATALYST-IRON
27
•»«
30
31
•*32
Catalyst
Burned
pyrite
40X
Reference
•
:A-28-ia8i"
10185
JA-35-3883*
!A-6l-5227h
;A-61-6823a
:A-53-13753f
Remarks
(Jrain size • 3.5 mm
uaia given ana apparatus ae-
under pressure
described •
necnanism or oxidation on oxide
catalysts discussed
Mechanically stable catalyst .
descrloed (consumption 2-3*/
month)
Apparatus for recovery of HjSO,,.
and fresh Pe^Oj from SD^A^
Fej03 Is described
Thermodynamics of Fe^O,, Ft,.
tlon compared.
o
z
(ft
z
H
O
TO
m
o
I
n
o
n
TJ
o
O
Z
-------
CATALYST - CHROMIUM
1
2
^
1;
5.
6
7 ' '
_8
Catalyst
rr-O-
Cr-O_
Cre°3
Cr3°n
2 3
Cr,0,
z 3
Crc°3
r>T- n
2 j
Cr-0,
Type of
Catalyse)
SoltH-
«as
n
n
n
ii
n .
n
n
Temp.
k<\a°c.
^50-6oc
50-550°
"Wi°r
6oo°c
"
550 "C
Weight %
Promoter
Rnn
^
'C SnO,,
D-10O?
; SnO-
3nOr,
T103, •
Al_0,
Weight %
Support
Material
•to _qo*
i? f? _
&t* J
nO TiO*
1H°T
Ti' J
"J.2u^
Weight %
Catalyst
'r.n -inf
£ 3
0-10051
;r2°^~J-1^*
broO^-12i
KnOn-Jif!
:r,on-io-
iitji
Porosity
Surface
Area
t
Gas
Flow
Rate
2OO-12O
Contact
Time
SO,
Cone.
7*
2-1UJS
7SS
o,
Cone.
J
2-l»
Conversion
Efficiency
50*
Ii.83.5*
55*
*f" 605f
80*
Reference
•
CA-53-21103e
CA-5^-211Olia
CA-SS-ailO^c
CA-59-2207K
CA-59-8l6la
CA-6l-1268lb
CA-62-1-1IO7C
CA-66-10867t:
1
Remarks
Theory of mixed catalysts was-
discussed. Maximum activity
did not coincide with max.
surface area.
Th» T-o«<-Mon- Rft in -
H
O
z
-------
o
z
o
m
8 M
> vo
a i-1
o
o
o
TO
T)
O
X
O
Z
10
11
12
13
IS
Catalyst
2 3
CHOH>2
Cr,0?
n
n
Cr~0. (I
n ^ "r\ / TI
Cr2o,
^g°S
Cri°n
Tn»oi
Catalyia
BBS
w
n
"
it
Temp.
I^U— *4DU'
515°C
n
Rin»c
Weight %
Promoter
Jase
L SnO,
n
"
Al 0
t 3
£
SnO-
^BaO aj^ri
Rp 0
2 3
H»n
Weight %
Support
Material
S10,
n
n
Waighf %
Catalyst
Cr nytrro-
oxlde hsdi
«•)
Pn H -
Surface
Area
G«s
Flow
Ran
vel . vo
_1OO
vel. vo
=11 5
yelj_yo
vel. vo
'=31-8*-
Contact
Time
.
.
SO,
Cone.
>6.»-8.<
°»
Cone.
[
Convernon
Efficiency
96.8-97.01
as*
qq 7
.
jA-so-geoi1*.
r*A_ai_7^A?
Remarks
of an active basic compound
t*V iDffl**t1f*n **' •"^»*«i»"*-
deBci>ib«d Ci> ffiar.AifMfc "inrp
resistant t« BAi«onin» f^hnn
the V-catalvsta.
and lab apparatus described.
CAfl M^QA Si.fl| 7.nA Al^A^
Bl_6... MnOp. Hid, CoQ '«^H
CuO 4ere f&und to have nesatlve
catalytic effect Mhen aAfi**
to Cr.j03-SnO, BaO and re£0.
active than baalc Sn-Or.
HjO, and Rcl comDared with
550°C troth CfttAlwwt «v*» Innmmo
to AS DOlfiAnl vty .
Pt^pAT^it 1 tf^p ff f d*.O P*1 rtPHr-
crlbed. c *
'
-------
CATALYST - CHROMIUM
17'
18
21
23
CrnatyO
ecgOa —
Cr-jO, or
rhrnmp nr
perlv. of
Cr
i*nnt A 1 nl n
Cr
tailing C
•nnri V
Cr,0,
ael
Cr,03
Type ol
Catalysis
Gas
"
n
ti-
ii
"
Temp.
Weight %
Promoter
BaCNO.^
Me.Cu Zr
Zn.Pb.Ag,
rare eart
t*l 1 o-ht 1 y
soluble
K^MA
Weight %
Support
Meter ial
Be, Cd,
Tl.
containln
K
Weight %
Catalyst
Containln
S10- and
cata.lv tic
active me
c nnin ni i nri 4
Porosity
contal
/•_j.y •
llv
Bl
Surface
Area
Inn
Gas
Flow
Rate
•
Contact
Time
SO,
' Cone.
°»
Cone.
Conversion
Efficiency
ID to not
Reference
•
CA-Sl-912ai
CA-22(P)-
-10183
-K865
I.
CA-35-3883"
CA-?5»P-
-•52 ^83
CA-lfl-3026'
Remarks
on enerKy of activation and
coefficient discussed. •
Oxidation nf - SO .e ntnl yii»H by
Fe.,0-.. Cr.O- ana coal asb.
boilers Investleated.
Mpf-.hnH nf p^ap^y.^^ ""^baiCt
mass desci-lbcd-
dBBCT*lbf»ri
ozlde catalyst ^}f)<;iissed
_
Contact Drooertleo of Cr^O.
dlRnmncrl 1 onimwintL m>
CA-tl-Z^fi^4 Is made.
•
O
(A
z
O
a
"i
w>
m
rv>
o
i
o
o
X
TJ
O
z
-------
CATALYST - CHROMIUM
24.
Catalyst
Cr~0,
Type of
CaMysii
Solid
gas
Tamp.
Weight %
Promoter
Weight %
Support
Materiel
Nl. Al, o
|flv_^lllea
Weight %
Catalyst
ea
i
Porosity
Surface
Area
Gas
Flow
Rate
Contact
Time
so,
. Cone.
°7
Cone.
Conversion
Efficiency
Reference
'-
CA-S^-lllSle
Remarks
effective Dare radius
and other adaomtlon
O
z
(/>
z
H
O
S!
m
VO
(JO
o
i
o
o
7)
•a
o
H
O
Z
-------
CATALYST - PIATINUM
1
•?
<4
5
6
7
Catalyst
Platinize
Pt
Hi
Pr
Pt
Pt
Pt-
Pt
Type of
Catalysis
1 Solid-
gas
11
"
ii
^
"
11
Temp.
lilO-^n"
•50O°C
20 °C
550-671
Weight %
Promoter
C
S10«
Weight %
Support
Material
platinize
asbestos
Weight %
Catalyst
ii
Porosity
Surface
Area
Gas
Flow
Rate
Contact
Time
SOi
Cone.
3.7-9*
°2
Cone.
Conversion
Efficiency
qq.6*
Reference
«
CA-Sl-l^la
CA-62-10061
CA^-imSe
CA-56-i0968c
CA-57-15865a
CA-55-7002f
Remarks
On Dlatlnized_ jnlcliroRie 76.756
S00 was oxidized.
"or the oxldn. of SO,,.
The possibility of electro-
chemical oxldn. of S0y by
electrolytic 0« was studied.
The activity of the PC catalyst
f^£creased when the eat&iygt
was -f-vley polarized, and
Increased when it was -vley
polarized.
Data Indicate that the actual
catalytic ac£ is the inter-
action of SO- with O2
adsorbed on the active
centers 01 trie cata-Lyst.
..
Kinetics of- the process SO_->-
1/2 0,^-S03 was studied in
the state of chemical eaullib-
riun on a Ft catalyst by the
aid of radioactive S.
Sorptlve catalytic process
or oxidation of SO,, on Pt
catalyst was studied.
.
o
z
(A
>
Z
-I
o
a
m
CO
m
o
o
•a
TJ
o
O
Z
-------
CATALYST -r PLATIMUH
a
9
10
1]
12
13
in
Catalyst
Platinum
(from
PtCl,,)
Platinum
Pt black
Pt
•
Pt
Pt
Pt
Typed
Catalysis
Solld-
Raa
"
It
n
IT
»
"
Temp.
uss-w
1|00°C
•?60-'iOO
350-475
>«7s»p
f
.
Weight %
Promoter
C
c •
C Al-0,
IP-IS*
Weight %
Support
Material
Silica Ge
Sllllmanl
bonded by
Et Silica
Asbestos
Si lira Oe
• Anhi»st:oR
Silica Re
SiO
88-qfiI
Pe-i°ii
Weight %
Catalyst
e«
e
5(t
0.024.
J.093*
Porosity
19-118
pore.dl
VJl
a
O
X
n
o
n
D
o
JO
O
z
-------
CATALYST - PLATINUM
16
17
»
Catalyst
Pt
Pt
•pt
Type of
Catalysis
Gas
n
n
it
Temp.
mo-fine
'•
*
Weigh! %
Promoter
.,
Weight %
Support
Material
Gel
Stlloa
Pellets (
Al,0,
1/8" Al O
Weight %
Catalyst
/2"1
3.2)1 Pt
Porosity
Surface
Area
•
Gas
Flow
Rate
•3500
1/l_hf
lllT.SlI
Ib/hr.
U&fta_Yn
lb/hr. (
of Bract
reac-tor
area) '
Contact
Time
•*
0. ft.
^
a. "ft.
•
so,
Cone.
o,
Cone.
Conversion
Efficiency
Bi» forp
treatment
76S
After Q2
• R»Tnr»
oxalic ac:
treatment
36. US
After 97.1
Reference
• «
fTwn pal'.^nrf]
)I
1
1
rA_iiii_7K37l
diffusion was dpt**i-ra1t^p^ t^y
measurlJiR the fate of
oxidation In a flow reactor
velocities.
n-,1,™ ^ «.~* ^ irmt-ajtlt
tci^j* 8^^*111 ^n* n f I11* nuu^^
Vftjlnelfcy wna- d*>«i*y1h«>i<9
o
z
(A
>
z
o
31
m
i/i
m
>
a
o
z
o
o
H
5
z
cr»
-------
CATALYST - PLATINUM
20
21
22
21
25
Catalyst
Ft
it
n
n
«
•
•t
n
n
n
11
Type of
Catalysis
Solid-
Gas
n
n
n
n
n
n
n
n
11
Temp.
360-130'
360-440'
127-S10
1)27-490'
400->)1S
350-450
7nn-4nn
450°C
460DC
Weight %
Promoter
c
c
p
c
c
c
c
p
Weight %
Support
Materiel
Silica ge
n it
Silica ge
• n n
Nlchrooe
(solral)
Weight %
Catalyst
0.2* Pt
0.5* Pt
Sofinfrv Pt
0.2 ran dl:
Pt wire
Pt net fr<
n nQ mm
Pt foil
n' 9 mm t-
0.05-4. OS
Rl (CT-ys-
0.1% Pt
Porosity
looX
ital
16500!
n
1i»p
1 plr
Surface
Area
30 so . i
71 80. I
0.17 nrj
(O.S-21
z 10 aq
H
It
Gas
Flow
Rate
/«.
/f.
m/2
m/g.
15-180
n
""IfoBO1
Contact
Time
b/hr
SO,
Cone.
?<
n
ti
n
n
n
11
n
°j
Cone.
(97J
air)
H
(t
It
n
ti
Conversion
Efficiency
f?iia-Jm<
Near 1001
Reference
•
CA-46-996lc
n
n
n
n
n
C A- U7.fi 2 "311 *
CA-66-982611
CA-2a-11DfiZ
CA-33-893^3
Remarks
10"K/per unit surface area-n.35
o n n it » -0;21
n n n n "«0.21
tt « ti R » .0.12
n n « n »» nQ.lt 8
n it n P ii «1.8o
Cpoelflo OQtolytia oottMtty
K YA^u*>A nalculat*»d fQl* ifaj^lr*11*
SAB, flow pates and df»tjr***e or
Use of Pt-olated Nlchrooe
noi»t-»>ir.ii. .rro^t- «r na< nh
bv AH.
-rro.t-. «k.._«_^
o
M
z
m
a
n
z
o
o
TO
•0
O
JO
O
z
-------
CATALYST-PLATINUM
3fi
27
28
30
31
32
Catalyst
p*-
n
Ft black
•
n
"
Type of
Catalysis
«?r»1 1 ri-
ll
n
ti
n
Temp.
^nn-fion
Weight %
Promoter
r.
Weight %
Support
Material
zel
Weight %
Catalyst
Pt Sp'nngp
Ft wire
0.04-0.06
thick
•Woven
Into net
Porosity
mm
Surface
Area
Gas
Flow
Rate
Contact
Time
so,
Cone.
°f
Cone.
Conversion
Efficiency
Reference
•
CA-Vl-71691
r«_lil_U1fi6r
CA-39-3719'
CA-S9-5160*
CA-141-2Q7S1
CA-47-1lBOe
rA_tio_p_7i «;«*
f
Remarks
Mechanism of SO. oxidation
Relation between reaction
order and enerzv of acti-
vation discussed for direct
Deslen of reactor zlven .
deacrlbed-'ualnK Pe.
M«r*hjan1 f*n af SO. ax1Hat-1nn
Rate data given for tvo Ft
catalvsts. different sizes.
Mpfhnrt nf preparing marrn
o
(A
z
o
m
a
o
o
o
TO
T>
O
X
>
H
O
•z
-------
CATALYST-PLATINUM
3C
38
Cetefytt
n
Type of
C*tatysi>
Gas
n
— • «
Temp.
Weight %
Promoter
Weight %
Support
Material
2 3
or SIC
Weight %
Catalyst
foil and
screen
Pt f»0.0
0.51)
Snnnirif Pt*
Porosity
11-
Suriaee
Area
f ^-Ilxl
_c_ra?/g
SponfEy
i 7*10
U|»f 10
20.6
Screen
22.6
Poll 6.
Gas
Flow.
Rate
f
t
?_
Contact
Time
so,
Cone.
°»
Cone.
Conversion
Efficiency
Reference
CA-i;?. 170551
CA-65-11396f
Remerkm
of Al-O- eel suDDort
d1neunRt«d
enerffv*?t KCnl/mnlf*.
ri4 B^llllH^rt
and reverae ^*^acfclon anri
enerav of activation
SO, followed bv means of
labelled 5
avatern
o
z
(ft
z
H
O
m
rn i—i
> vo
o
I
o
o
TJ
O
TO
H
O
Z
o
-------
CATALYST , PLATINUM
HO
13
11
Catalyst
Pf
11
II
"
Type of
Catalysis
^nl 1 A
"
n
"
Temp.
1(50-800
•1100-15
Weight %
Promoter
'C
)°C
Mn 5*
for Fer
Ml Al •
r.u R1
Mn RA V
Sn, CrJ
Weight %
Support
Material
Silica
Gel
Clay con-
taining
Al,0-.-SiC
Silica xe
Weight %
Catalyst
Platinize
ntchrome
wire
solrals .
•Platlnlz
platinum
Porosity
-
id
Surface
Area
Gas
Flow
Rate
15n3/hr
•68-15C
Contact
Time
SO,
Cone.
66%
•7*
o,
Cone.
3«»
Conversion
Efficiency
60J
•69.1-76.'
. X
Reference
CA-IH-P-SUlt
rA_";p_ifiRc;ia
CA-28-P590
Remarks
with Ca_V anri I?o_^^-V
catalysts
Method of Dreoarlnn SO
deacrlbtid
size and oorosity and the
calculation of overheating
£
BtiHl) on ^-^^ ^At.AlyMr*
oxidation of SO^ WAR
Investigated. ^i*93fiui-<>
and soace velocity on
activity of Platinized
Nlchrome and nlatlnun
lnvpRt-.1gat:pri
=
Methods of preparing rar.nly«f«
AV*A CTlvtAM
2
O
(A
>
z
H
O
m
M ro
m o
o
i
o
o
X
13
O
O
Z
-------
CATALYST-PLATINUM
«7
K8
19
50
Catalyst
Silicates
of Pt OP
Pb HT'Oun
metals
Pt
**
II
Type of
Catalyst
rtnn
n
R
n
n
Temp.
Weight %
Promoter
Al Tl
Per Cu.
Zr. Zn.
Pb. AK.
Ce, Nl.
Co . B and
rare
pp**rhM
Weight %
Support
Material
Al RA
rrt 7r
?n or Tl
and S1CU
"
KleaelKuh
Dunlee or
faucoall
Weight %
Catalyst
,
•
Porosity
Surface
Area
Gas
Flow
Rate
Contact
Time
SO,
Cone.
°l
Cone.
Conversion
Efficiency
Reference
•
CA-26(P)-202;
CA-11-1317S
SA-lS-lfla^"
:A_nn_ponB
Remarks
Oatalla nf gafalytt pnepajatlon
Pd YfiP-
r
Methods of catalyst
V.catalvst
M«
-------
CATALYST - PLATINUM
53
53
55
56
57
Catalyst
"
»
11
Ft com-
pounds
Pt
Type of
Catalysis
Oas
n
"
II
II
n
Temp.
iai^525!
•U16-U6<
125°C
Weight %
Promoter
"C
Fe com-
pounds
Weight %
Support
Material
Asbestos
Silica ge:
Asbestos
Weight %
Catalyst
Porosity
Surface
Area
Gas
Flow
Rate
Contact
Time
SO,
Cone.
\
•81
•9«
•0.3-
0.7»
•
Cone.
• I?*
Conversion
Efficiency
•50-98*
•Ql-97*
Reference
CA-19-2111
CA-21-29623
CA-2Q-P-S610
ra_3li D firrt«
Remarks
activity discussed.
Two types of converters
Influence of reaction rate on
ope. ratlnir conditions in
Hactrylhori
_
f^0"!^",1"^1^" Pthmllt^-et
M
o
z
M
Z
o
C!
ro ro
z
o
o
O
Z
-------
CATALYST- PLATINUM
59
60
62
Catalyst
Pt
it
n
(Electric
ally char
Trpeof
Caulysa
Solid-
gas
n
n
;et And
rnfnlyMr arflylty lmrent'1*
pn» rharga rtncraaiog — _— ^~
tlon described
o
I
H
O
C!
ni ro
> o
jjj u,
o
o
TO
TJ
O
H
5
z
-------
CATALYST » PLATINUM
65
•67
•69
•70
Catalyst
Pt
•-„ — •
Pt-Pd
7'5*-21;*^
Pt-black
Type of
Catalysis
Solid-
fl
II
fl
Temp.
1150.650'
U89°C
160-603°
Weight %
Promoter
:
1oO'
l.SsU.Q*
:
Weight %
Support
Materiel
SID.,
Phnr^Ofll
iilica gel
Lsbestos
>r calcine
igsn.
ft 1 4 to
si 1 1 na ft
AUO
nv* Mp^rt
"~ *4
Asbestos
Weight %
Catalyst
3%
1.5*
J
4
0.6*
7.5*
Porosity
"*°5 —
• r .
Surface
Area
Gas
Flow
Rate
6000 v<
Contact
Time
L.
SO,
Cone.
331
8*
"I
°i
Cone.
17*
60*
Conversion
Efficiency
llT.Q-Qfi :
I
9°*
97.2*
Rn i-9< p
16.2-91.1
ft
Reference
«
CA.6B-72717r
RT-^fl_155R
CA-42-P-2110
PA 3n_^°
Remarks
Examination of structure of
bv' electron mlacroaconv
dnftf nlhf^ri .
Kechanlam Inveatmatea.
Klnrri ,rMn,«
Pt /*nmpnT>»ri w1 ^h TQ P'1'1 Prl
Catalvat nrepat>at1nn rt»s_
r.l-lbArt
Platinua blaak aanporod
«: j £
o*J{lat. 1nn enmpa**»>fi
o
z
01
z
m
V) IY>
m o
o
X
o
o
TO
T)
O
-------
CATALYST - CARBON
1
2
3
D
5
b
7
Catalyst
ferbon
Carbon
larbon
Active
cn&r .
:oal
Carbon
Activated
Charcoal
Type of
Catalysis
Solid-
it
11
n
Tamp.
20 "C
65-80 °C
20 °C
Weight %
Promoter
Na.CO.
Weight %
Support
Material
WfOIQnt «D
Catalyst
1
Surface
Area
571-
1289D^
!65-3'(2
m2/s
Gas
Flow
Rate
Contact
Time
10-20se<
SO,
Cone.
b-3*
°?
Cone.
Conversion
Efficiency
30*
7
Reference
•
:A-62-P-1357f
:A-62-8U2Sc
:A-6^-?662h
:A-63-8071e
see EI-66-
3001 card)
IA-65-P-
iRaosR
IA-65- 193510
CA-17-3Z718
Remarks
Poor conversion of SO« was
catalyst retain H^SOi -
The mechanism of catalytic
ixldn. on activated carbon.
The contribution of carbon free
radicals In the SO- oxldn. at
1th continuous or Interm^t^ept
.downwashiruj. Dll Hr-SOj, could
be recovered. " "*
Bcrubblna tower fqf nmY^Ul
SOj. from the flue oafiea bv wet
catalytic oxldn. to H— SOi
over coal are described.
..
Correlation between the concen.
of free radicals of carbon and
Its catalytic activity In H~S
and SOn oxldn. processes was
studied.
. coal aa to aetlvltv noted.
• Quantitative meaaureoenta •
^Iven.
•
o
(A
z.
H
O
•a
m
u>
m
o
X
n
o
x
TJ
O
X
O
Z
ro
o
ui
-------
CATALYST-CARBON
B
9
10
11
12
13
It
Catalyst
Carbon
Carbon
Carbon
Carbon
Carbon
Carbon
Activate!
coal or
charcoal-
Type of
Catalysis
Solld-
Oas
Solid—
Qas
Solld-
Oaa
Solld-
uas
Solld-
oas
Solld-
Qas
Solld-
.uaa
Temp.
20°C
yn-Ticf
20°C
<220°C
o-eo^c
-
.
Weight *
Promoter
Sulfur
Weight %
Support
Material
'
Weigh,*
Catalyst
Porosity
'ore '
•adlus
<20A
Surface
Area
>52-B92
1/r/3
Gas
Flow
Rate
"Linear
rate -
15-20
cm/sec
Contact
Time
SO,
Cone.
1-5*
L0.25»-
o,
Cone.
,
»*
Conversion
Efficiency
^75-92*
Reference
•
CA-^^-171fi7^
ri_K/;_liflr.
CA-59-67C
CA-51-B391h
CA-56-10963e.
!A-61-275Bc
IA-61-P-3961C
Remarks.
Catalytic oxidation of SO, In
presence of HgO on activated •
C surface discussed. Good
catalyst on.ly In llq. cont. '
method.
Mechanism of. oxidation at car-
Don surface discussed.
Relation between oxidation and
the adsorption and desorptlon
Isotherms on activated carbon
studied.
Rate of oxidation depends on
pressure.
Effect of temperature on con-
version In liquid-contact
method.
Various grades .of carbon and
metbods pr activation for ad-
compared. Influence of cone..
particle size and surface area
Investigated.
Anthracite coal Rives better
conversion than charcoal or
peat coal .
o
z
u>
z
o
m
o
x
o
o
3J
3
X
-\
o
z
ro
o
en
-------
CATALYST-CARBON
U>
z
H
O
m
M
ra fu
> o
IP
o
o
X
3
Weight %
Promoter | Support | Catalyst
Material
Surtaea I Gas I Contact I SO,
-------
CATALYST-MANGANESE
1-
2.
3.
Cetatrst
KMnO,, ant
MnSAfi
KnO]
Type of
Catalysis
Gas
Solld-
Qaa
Solld-
Oas
Temp.
Weight %
Promoter
Weight %
Support
Material
Weight %
Catalyst
,
.
Porosity
-
.
Surface
Area
Gaa
Flow
Rate
v
H:-
Contact
Time
.
SOi
Cone.
•
«2
Cone.
Conversion'
Efficiency
90S
Reference
m
cA-es-ioea
CA-20-P-
-57862
IA-67-P-
-57063k
Remarks
movlnu S comoounds from [cases'
reviewed .
• .
Methcu^ or pAtalvat nTTP^rA'
tlon described
.
o
M
z
o
O
X
o
o
X
TJ
o
H
O
z
-------
MISCELLANEOUS SOLID CATALYSTS
2
3
4
>5
U
Catalyst
Pt-AU
alloys
Ft. V^C..-,
Fe?°3~ '
Com. Na-
vanadate
Pr. qpnn^p
VoO£ &
FB,8,
Tnn_PYf*ha
Ft or
v n
*• 3
Type of
Catalysis
gas
II
"
It
,£0 "
11
Temp.
580 "c
feo-65°C
U50°C
Weight %
Promoter
Weight «
Support
Material
Vlnvl-
yrldlne
Weight %
Catalyst
( 0 . 1'mm
dla) or
pi AtlPfi
Porosity
Surface
Area
Gas
Flow
Rate
loiJ/min
Contact
Time
3hrs
10 mln
l.ll sec
SO,
Cone.
1-1. 5*
0.15-
2.3l
0.13*
°*
Cone.
20.8^
0-6.02*
Conversion
Efficiency
50*
-
70*
ytf,
Reference
•
CA-55-21t99h
CA-56-15395K
CA-56-12355a
CA-62-P-2525<
CA-63-P-163'
Remarks
Ft. Au,,. Pt-Au alloys. Cr.
Rh. Ae'was evaluated.
Isotherms Tor the catalytic
the abs. soeed of reaction a,s
a function of equil". state is
examd. for 3 different cata-
lysts, Pt, V00C and Fe00,.
Fore structure and Doff of
different catalyst specimens
were detd. und^r different
terns, conditions, and effects
on reaction kinetics Cfilctd.
Catalytic oxidn. of SO- as a
function of the residence tir.ie
of the gases in the reaction
chamber was studied.
Ion exchange resins were used
to oxidize SOo In Industrial
gases to SO,.
i Absorption capacity of C-
contg. absorbants for S-concg.
gases can be greatly in-
S0n to SO .
c J
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-------
MISCELLANEOUS SOLID CATALXSTS
7
9
10
12
13
14
Cits**
Chromium-
'anadlum
Pd
Pt-Pd
Zno
5no2
Rh-black
tteSOi,
Type of
Catalysis
Solid-
Gas
Caa
Solid-
Gas
Solid- '
Has
uas
Solid-
las
Solid-
Gas
Solid-
Gas
Temp.
150°C
150-500'
U70-500"
Weight %
Promoter
•
* i
Weight %
Support
Material
Asb**s toa
Silica Qe:
u,o,.Mgs(
illloa Qe!
>r asbestc
:iay
Asbestos
Fibers
Weight %
Catalyst
0.05-2-U
u*0.3-0.6j
''lO-lUOJ
3
u.ubmm
Porosity
*t
tnicK
Surface
Area
'Gas
Flow
Rate
'ol.Vel.
1 300
Contact
vel.=
78.3
8-50 I/
iQOO vo]
foi cac^
Contact
Time
ir
t
\r
SO,
Cone.
18%
o»
Cone.
\
Conversion
Efficiency
95.3*
78. 8-97. 2)
Reference
«
CA-32-722o»
CA-117-523* 1
CA-lll-P-
-4625e
JA-IT-P-SSTS'
:A-20-5»
:A-42-3918d
Remarks
Netriod or preparing cataxyst
descrioea. energy 01 activation
Klven.
Pt-blacx.
icas flow rates and degree of
conversion. Compared witn pt.
Advantage of Pt catalyst con-
crystalline catalytic structures
described.
as catalyst for .gaseous oxida-
tion and reduction reactions.
Data summarized on oar tic IP size
S porosity t the calc. of over-
neating coerrici.ents for several
catalysts used for the 'oxidation
or bu?.
.
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to
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-------
MISCELLANEOUS SOLID CATALYSTS
15
if.
17
IB
20
Catalya
Ch,Tl-ani
Ta ozldei
from
Loparite
Rf
n-tyoe
semicon-
ductors
(HO^.Pe,!
P102.As20
v,o5)
Deriv. o:
CO.CU.W,'
U.or Ho
Cr.Mn.As
Sb.Ta.Nt)
or ,DI
V.Ho.W.U
Cr ,Mn,Tl
01 ri.v,
Pe or Nn
Type of
Catalysis
Solld-
Oaa
Cnl 4 ft
Qas
s.
,
Solid-
das
Solid-
Gas
Qas
Tamp.
550°C
•H50-60C
"C
Weight %
Protnolar
Zeolite
Metals
Zeolite
'contain-
ing K
Weight %
Support
Material
Zeolite
contalnlr
Al.Be.Cd
£r f&nf 11
r^pif^lgiihr
».»n1-1t:
•z.
-\
o
X
m
<2
m
o
X
o
o
X
-o
o
o
z
ro
-------
MISCELLANEOUS SOLID CATALYSTS
22
23
24
25
26
27
Catalyst
Mixture
of CbOi;,
Ta?Oc; anc
TlOo froo
Loparlte
Pt V,0q+
KzSoii.1
Na,$0A,
S10,,CrO-
Pe.j0j.cuc
ZnO.NIO
and V2O5
CaO,
Ca(OH)3.
CaCOj .
CaHCO-i.o:
similar
Mg.Sr.Ba
eompminrtf
Al-O,
Cr oxide
Cu oxide
and/or
Mn oxide
Type of
CflatysB
Solid-
Gas
Solld-
Qas
,
Solld-
Oas
Solid-
Gas
solid-
Gas
Solid-
Gas
Temp.
ISO-SSI
°c
1st -
400-501
2nd -
500-HOl
Weight %
Pfomotei
0-2SS
NaOH.
Nago,
Na2CO,,
KOH.K^O,
K,CO?,
or KHCOj
•c
•c
Weight %
Support
Material
Al-O-j.
S10?-Al7(
or dlator
aceous
earth
Silicate:
Weight %
Catalyu
1
-
971
Porosity
Surface
Area
,
Gas
Flow
Rate
!SO-250C
iJ/mJ
iatalvst
'hr
Contact
Time
SO,
Cone.
5.1*
°»
Cone.
«b.5*
»
Conversion
Efficiency
teo*
Reference
*
CA-32-22953
rA-SS-SBBS1*
CA-57-1101
CA-64-(P)-
6370d
CA-66-P-
59305U
CA-68-P-
93771v
Remarks
Use or this catalyst in contact
process aescrioes.
Mechanism of oxidation on oxide
catalysts discussed.
Mechanism discussed.
from coke oven gases and
method or cataiysi preparation
described.
Regeneration of activity dis-
cussed.
tlon or exhaust teases containing
s-compounds described.
s.
o
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m
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CATALYST-IRON (Liquid-Phase)
1.
2.
3
Catalya
Vfsa^
fe salts
Type of
CataJyaa
I.lqii120-90°(
•
Weight %
Promoter
Weight*
Support
Material
Water
Weight *
Catalyst
•3-20S
PeSOn
Porosity
Surface
Area
Gas
Flow
Rate
lO-l./hi
Contact
Time
SO,
Cone.
•1»
°»
Cone.
'3-15*
COftV6TB*On
Efficiency
Reftfencv
•
IA-50-1509h
-6992h
"A-P7-SBQO1*
Remarks
fhrlrte»i--4rt»i nf SO- In ""lutlQP
of li«*m aulfate descplb^lf
HoMKonlcn*. «14 on..B«>Arf '
'
bv ll 1. ft Ao.1. V&A O In Mnnnnnfe j*||f: 0fl
«n1u4:1nnA nr bjktilf* Al^ffSAij)
ril R«**ia*iB*(f .
Cfttfll^tlc oxldatloi^ ^y ^on—
Ized solutions of heavy Qf£^Bls
dcscfibod
-
-
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CATALYST-NAMCANESE (Liquid-Phase)
1
2
3
II
5
6
7 '
8
9
10
.
Catalyst
nnsoj.
solutions
HnO Nn(OH
KMnOa
solutions
HnSOii
solutions
Hn,(SOi,)j
solution
Hn
catalyst
MnSOi|
MnSOi,
Mn
catalysts
MnSOt,
HnSOii
-
Type of
Catalysis
Liquid .
oas
, Llaul
3 Gas
Liquid-
Gas
Liquid-
Gas
Llauld-
Qas
Liould-
Oas
Llquld-
Cas
Liquid-
Gas •
Liquid-
Liquid-.
Gas
Temp.
_
0-20°C
0-60°C
20-«0°C
Weight %
Promoter
Weight %
Support
Meterial
Weight %
Catalyst
Porosity
Surface
Area •
Gas
Flow
Rate
Contact
Time
SO,
Cone.
o,
Cone.
Conversion
Efficiency
Reference
-
EiU9&5JZ1b5
CA-^0-20908
CA-3D-2M7M-'
CA-10-15962 -
CA-30-3596"
CA-30-57323 .
CA-10-70195
CA-31-19633
CA-32-52859
CA-ie-zzo/a
Remarks
Rate of reaction of SOp and
02 in aqueous nnso^ solution.
Investigated.
Mechanism of oxidation of SO?
In solution of .Hn salts. Oxl-
KlnPtlcR dln^URAPd,.
Effect of gas flow rates and
poisoning of Hn by phenol.
Effect of temperature on
Effect ot poisoning ln oxlda_
tlon of. 503 by.,ozone.
H?SOu discussed.- • -
Mechanism of oxidation of S02
lyst studied by ootlcal means.
Composition of ncynhhei. llqi.lrt
In NH, scrubbing of flue aaa
catalyst.
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CATALYST-MANGANESE (Liquid-Phase)
11
12
13
ID
15
* •
Cetstrst
Hn oxides
KnSOh
MnSOfe
MnSO,.
solution
contain-
ing Al
MnSOj,
solution
contain-
inff Al;(!
Tvpeof
****
Llquld-
aas
Llquld-
uao
Llquld-
uas
Liquid-
Gas
Liquid-
Gae
lh i?
Temp.
15-50-C
25°C
Weight «
Promoter
0.15* A]
Weight %
Support
Material
Weight*
Catalyst
01
0.05-
l.U»
0.051 Mn
0.05-5*
Al
0.05* Mn
Porotitv
Surface
Area
Gas
Flow
Rate
Contact
Time
...
SO,
Cone.
0.17*
1-2O«
o.
Cone.
2.8*
C^^ior,
Efficiencv
65*
Reference
•
:A-6l-P-6657e
JA-63- 17193d
;A-34-P-38b9'
3A-'t7-31l6b
Remarks
necnoa 01 H^SUJI proaucnon xroc
flue Rases described.
Concentration of NnSOn and re-
J^fr^^r. nf HnSOi] eoneen^r^t^on
on catalytic activity of MnSO*
in oxidation or &t>2 by 83.
Method of preparing 40* ffoSOa
described.
Air oxiaation auring scrubbing
with aqueous solution In a
column Is ciescriDea.
•
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-------
MISCELLANEOUS LIQUID PHASE CATALYSTS
1
2
3
It
*5
CatalyM
*.i». of
Mn
Aqueous
solution
of Al and
In com-
pounds
Metallic
catalysts
!IHi
Iodine o
metals (H
UU , ftU , ^
Tl , Fe , '&
HI, CO, S
AS, CM, V
Mo
Type of
Cstalyiis
Tin
Llauld
Liauld
Liquid
Llauld
,
,
,
,
Temp.
<100°C
Weight %
Promoter
i
Weight %
Support
Material
0.05X Mn
0.05-5*
Al
Carbona-
n»Qnfl Ad-
sorbents .
Meerschau
nnrt Alnrnl
nun Slli-
r^t:ea
• •
Weight %
Catalyst
i
Porosity
Surface
Area
Gas
Flow
Rate
\
2000m3/
Contact
Time
r
SO,
Cone.
fl-20*
Us/m3
.
0,
Cone.
Conversion
Efficiency
.
Reference
«
:A-27-S8q9l
3A-3f-(P)-
38B91
CA-65-CP)-
5210h
CA-Ot>-lOOO94l
CA-6a-P-8706(
Remarks
Catalytic oxidation ey ionized
solutions of heavy metals ae-'
Method of preparing 401 HpSOi,
descrioea.
Gas la scrubbed with water and
catalyst.
Kinetics of SOo-NH^-llquld
HjO system discussed.
SOp was catalytlcally oxidised
to so-% wnicn was aosoroed By
dilute H,SO,,.
.,
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z
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8!
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o
o
a
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-------
MISCELLANEOUS QAS PHASE CATALZSTS
1
2
3
H
5
6
7
Citalyct
Oxides oi
H, 1
NpO In
the pres-
ence of
HO
Oxides o
N2
Typed
CstBlytB
etas— aai
.Iduid-
;as
gas-gas
tas-eas
caa-aa:
Temp.
80-130«
4QO°C
Weight %
Promoter
Weight %
Support
Material
Weight %
Catalyst
Eaulvaler
to It KNC
Porosity
1
'•
Surface
Area
Gaa
Flow
Rate
lOem/sc
Contact
Time
SO,
Cone.
4-161
0.31
;
o,
Cone.
0.2-
15- BJ
Conversion
Efficiency
67-92*
Reference
•
CA-49-J)24Bi
IA-57-10570C
CA-60-21DC
!A-61-327b
CA-6H-173358
:A-65-17756f
IA-65-506lf
Remarks
The effect of temp, and "26^
concen. on the conversion or '
SO^ to 3U% was scuaiea TAP biax .
liquid spray rates. '
Decomposition of N,0 at 700°C >
can be used as a source of 0
atoms x~or me A^iaadon oT Vuj.
The effect of concen. of R?0j
and H2S04 In nitrous aclda on
the conversion of SO? ln tne
Kachkaroff process is stuaiea.
Oxidation of SO? with nltroae
In packed columns and In sieve
plate columns was studied.
Air concentration 0-3t SO? Is
converted to (NHi, )2SO» .»» »»?
and 0-j.
Reactions NO,+SO, *JIO+SO, and
NO^+so?^CNO2+so^ vary witn" tne
concentrations of the reaetants
and with temperature.
Absorption rate of SO? la the
presence of the N oxides was
Increased by approximately
20-51 In comparison with the
absorption rate In the absence
of N oxides.
.
O
(A
O
d
m
x
o
z
o
o
O
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ro
-------
MISCELLANEOUS GAS PHASE CATALYSTS
8
9
Catalyst
H3°3
N oxides
Type of
Catalysis
KILS— K&.S
gas-gas
Temp.
700°C
Weight %
Promoter
Weight %
Support
Material
Weight %
Catalyst
Porosity
-
Surface
Area
-
Gas
Flow
Rale
•
Contact
Time
SO,
Cone.
Cone.
Conversion
Efficiency
1
Reference
CA-53-13753
CA-t>7-23705«
~
Remarks
Thennodvruunlr-a of NiO.jr Ft. V->O
and Pe,0? catalytic oxidation-
compared .
Factors 'affecting the formation
or 503 in flame gases discussed
z
o
a
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(A
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o
TO
TJ
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X
H
5
-------
APPENDIX IV
CAPITAL COST SUMMARY SHEETS
219
• MONSANTO RESEARCH CORPORATION •
-------
APPENDIX IV
INDEX
PROCESS NAME Page
Monsanto-Penelec 221
(low temperature effluent)
Monsanto-Penelec
(high-temperature effluent) 222
Klyoura-T.I.T.
(low temperature effluent) 223
Klyoura-T.I.T.
(high temperature effluent) 224
Reln-luft Process
(low temperature effluent) 225
Sulfacld Process
(low temperature effluent) 226
Mitsubishi Process
(low temperature effluent) 22?
T.V.A. Process
(low temperature effluent) 228
Gallery Process
(low temperature effluent) 229
220
• MONSANTO RESEARCH CORPORATION •
-------
SO2 REMOVAL FROM FLUE GAS
Category: Existing Power Plant (Low Temperature Effluent)
Capital Cost Estimate Summary
Name of Process: Monsanto-Penelec Flue Oas Rate: .2.5 MMSCPM
MW 1400
Cost $
1) Purchased Equipment
Flue Gas Heater 655fQQQ
Catalytic Converter
Primary Heat Exchanger 1
Second Haat Exchanger 2
Mist Eliminatof 1,850,000
TOTAL 7.175.000
2) Fixed Capital Cost* 34,081,000
3) Working Capital** 3,408,000
TOTAL INVESTMENT 37.489,000
Capital Requirements
$/kW Capacity 26.77
* Langs factor applied; 4.74 for fluid process plants
**102 of the fixed capital cost
221
• MONSANTO RESEARCH CORPORATION •
-------
SO2 REMOVAL PROM FLUE QAS
Category: New Power Plant (High Temperature Effluent)
Capital Cost Estimate Summary
Name of Process: Monsanto^Penelec Flue Oas Rate: 2.5 MMSCFM
MW 1400 .
1) Purchased Equipment
Catalytic Converter
Economizer
Air Preheater
Mist Eliminator
Cost $
540.000
1.530.000
2.600.000
1.850.000
TOTAL
2) Fixed Capital Cost*
3) Working Capital**
TOTAL INVESTMENT
4) Capital Requirements
$/kW Capacity
6,520,000
30,970,000
3.097.000
34.067.000
24.33
* Langs factor applied; 4.74 for fluid process plants
*»10* of the fixed capital cost
222
• MONSANTO RESEARCH CORPORATION •
-------
SO 2 REMOVAL PROM FLUE QAS
Category: Existing Power Plant. (Low Temperature Effluent)
Capital Cost Estimate Summary
Name of Process: Kiyoura-T.I.T. _ Flue Oas Rate: 2.5 _ MMSCFM
moo
Cost $
1) Purchased Equipment
Flue Qas Heater 655,000
Converters, Blowers, and Motors 5*10.000
Primary Heat Exchanger 1,530,000
Secondary Heat Exchanger 1.0*40.000
Electrostatic Precipitator 1.200. OOP
TOTAL 4,965,000
2) Fixed Capital Cost* 23,535,000
3) Working Capital** 2.354.000
TOTAL INVESTMENT 25,889,000
Capital Requirements
$/kW Capacity 18.49
* Langs factor applied; 4.74 for fluid process plants
**10JI of the fixed capital cost
223
• MONSANTO RESEARCH CORPORATION •
-------
SO2 REMOVAL FROM FLUE GAS
Category: New Power Plant (High Temperature Effluent)
Capital Cost Estimate Summary
Name of Process: Klyoura-T.I.T. Flue Gas Rate: 2.5 MMSCFM
MW 1400
1) Purchased Equipment
Convertersf Blowers, and Motors
Economizer
Air. Preheater
Electrostatic Precipltator
Cost $
i 530,000
1.200.000
TOTAL
2) Fixed Capital Cost*
3) Working Capital""
TOTAL INVESTMENT
4) Capital Requirements
$/kW Capacity
4.31Q.QQQ
2G.ino.ono
2,043,000
22,473^000
16.05
" Langs factor applied; 4.74 for fluid process plants
of the fixed capital cost
. . 224
• MONSANTO RESEARCH CORPORATION •
-------
SO2 REMOVAL PROM FLUE GAS
Category: Low Temperature Effluent
Capital Cost Estimate Summary
Name of Process: Relnluft Process Flue Qas Rate:
MW 1400
2.5
MMSCPM
1) Purchased Equipment
Adsorber
Regenerator
Heater
Cooler
lowera and
Cost $
inn .nnn
110jOOO
3,936,000
TOTAL
2) Fixed Capital Cost*
Sulfurlc Acid Plant
TOTAL
3) Working Capital**
TOTAL INVESTMENT
4) Capital Requirements
$/kW Capacity
5,441,000
25,800.000
2.600.000
28.60Q.000
2.860,000
31.^60.000
22.47
* Langs factor applied; 4.74 for fluid process plants
of the fixed capital cost
225
• MONSANTO RESEARCH CORPORATION 0
-------
SO 2 REMOVAL FROM FLUE GAS
Category: Low Temperature Effluent
Capital Cost Estimate Summary
Name of Process: Sulfacld Process Flue Gas Rate: 2.5 _ MMSCFM
MW
Cos.t $
1) Purchased Equipment
Venturi Scrubber
Reactors _ 7,700.000
Blowers and Motors 810,000
Acid Purification and Concentration 1,175.000
TOTAL 13,025,000
2) Fixed Capital Cost* 61,750,000
3) Working Capital** 6,175,000
TOTAL INVESTMENT 67,92$,000
Capital Requirements
$/kW Capacity
» Langs factor applied; t.7l for fluid process plants
••10* of the fixed capital cost
226
• MONSANTO RESEARCH CORPORATION •
-------
SO2 REMOVAL FROM FLUE GAS
Category: Low Temperature Effluent
Capital Cost Estimate Summary
Name of Process: Mitsubishi Process Flue Qas Rate: g.q MMSCFM
MW IHOQ
Cost $
1) Purchased Equipment
Absorber 525,000
Cyclone Separator 791,000
Electrostatic Preclpltator 1,200,000
Ammonia Scrubber 30,000
Oxidizing Tower 30,000
Air Compressor 170,000
Crystallizing Equipments 1«MO»000
Miscellaneous 200.00
TOTAL JJ.386. OOP
$/kW Capacity
2) Fixed Capital Cost- 20,790,000
3) Working Capital" 2,079,000
TOTAL INVESTMENT 22,869,000
Capital Requirements
16.33
• Langs factor applied; b.Jb for fluid process plants
••10JR of the fixed capital cost
22?
• MONSANTO RESEARCH CORPORATION •
-------
SO2 REMOVAL PROM FLUE GAS
Category: Low Temperature Effluent
Capital Cost Estimate Summary
Name of Process: T.V.A.-Sulfurlc Flue Gas Rate: p.5 MMSCFM
Acid Process
MW
Cost $
1) Purchased Equipment
Evaporator 1.450.000
Packed Scrubber 7.700.OOP
Blowers & Motors 810.000
Mist Eliminators 1.850.OOP
Capital Requirements
$/kW Capacity
TOTAL 11
2) Fixed Capital Cost* c;£ nnn nnn
3) Working Capital** 5,600,000
TOTAL INVESTMENT 61,600,000
* Langs factor applied; 4.74 for fluid process plants
**10$ of the fixed capital cost
228
• MONSANTO RESEARCH CORPORATION •
-------
SO2 REMOVAL FROM FLUE GAS
Category: Existing Power Plant (Low Temperature Effluent)
Capital Cost Estimate Summary
Name of Process: Gallery Chemical Flue oas Rate: 2.5 MMSCFM
a -Process "
MW
Cost $
1) Purchased Equipment
Flue Gas Heat Exchanger 2,140,000
Catalytic Reactor 300.000
Absorber-Stripper 220.000
Furnace ^QQrQQQ
Gas Cooler 65.000
Sulfurlc Acid Absorber & 40.000
wist Eliminator
Acid Cooler 100.000
Fan and Blower 63.000
TOTAL 3.228.000
4) Capital Requirements
$/kW Capacity
2) Fixed Capital Cost* 15,300,000
3) Working Capital** lr»53QrOQQ
TOTAL INVESTMENT 16,830,000
* Langs factor applied; 4.7*4 for fluid process plants
**10$ of the fixed capital cost
229
• MONSANTO RESEARCH CORPORATION •
------- |