PB   198  809

APPLICABILITY OF  CATALYTIC  OXIDATION  TO THE
DEVELOPMENT  OF  NEW  PROCESSES  FOR  REMOVING
S02  FROM FLUE  GASES  -  VOLUME  II --EXPERIMENTA L
PROGRAM

R.  E.  Opferkuch,  et al

January  1971
         NATIONAL TECHNICAL INFORMATION SERVICE
                                                Distributed .,. 'to foster, serve
                                                   and promote  the nation's
                                                      economic development
                                                          and technological
                                                            advancement.'
                                                U.S. DEPARTMENT OF COMMERCE
                 This document has been approved lor public release and sale.

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  APPLICABILITY OF CATALYTIC OXIDATION

TO THE DEVELOPMENT  OF NEW PROCESSES

  FOR REMOVING  S02 FROM  FLUE GASES

    Volumell-  EXPERIMENTAL PROGRAM


            Contract No. PH 22-68-12

                  Prepared by
                 R.E. Opferkuch
                  S.M. Mehta
                 M.G. Konicek
                  D.L. Zanders
                   Submitted to

         . Division of Process Control Engineering
        National Air Pollution Control Administration
            Environmental Health Services
              U.S. Public Health  Service
      U.S. Department of Health. Education, and Welfare
                5710 Wooster Pike
               Cincinnati, Ohio 45277

-------
 BIBLIOGRAPHIC DATA
 SHEET
                  ). Report No.
                    APTD-0676
                                           3. Recipient's Accession No.
                                                             5. Report Date
                                                              January 1971
4. Title and Subtitle
  Applicability
of Catalytic Oxidation  to the
  ment of New  Processes for  Removing  SO  From
  Gases     Volume II  -  Experimental Program
Develop
Flue
                                                            6.
7. Author(a)IT E.  Opferkuch (Project Leader)
 S. M.  Mehta,  M.  G. Konicek,  D.  L. Zanders
                                                            8. Performing Organization Kept.
                                                              No.
>. Performing Organization Name and Address
 Monsanto Research Corporation
 Dayton  Laboratory
 Dayton,  Ohio   45407
                                                            10. Project/Task/Work Unit N-
                                                            11. Contract/Grant No.
                                                               PH 22-68-12
11 Sponsoring Organization Name and Address
 Process Control  Engineering Program
 National Air  Pollution  Control  Administration
 Environmental  Health  Service,  U.S. Dept.  of HEW
 5710  Wooster  Pike
 Cincinnati, Ohio  45277
                                                            13. Type of Report 4t Period
                                                               Covered
                                                            14.
IS. Supplementary Notes
16. Abstracts
  •During  the first phase  of this  project, which covered  the  identification
  and evaluation  of existing and  potential  methods  of applying catalysis
  to the  oxidation and removal of  S0_ from  power  plant  stack gas,  a large
  quantity  of information  and data was accumulated  and  assessed.   This
  evaluation revealed the  need for laboratory verification  of the  publishe|d
  data on  promising oxidation systems.   This volume presents subsequent
  data and  information generated  in the  laboratory  and  on  the drawing
  board.
17. Key Words and Document Analysis.  17o. Descriptors
  Tests
  Catalysis
  Oxidation
  Catalytic  Converters
                                                   Sulfur  dioxide
                                                   Flue gases
                                                   Electric  power  plants
I7b, Identifiers/Open-Ended Terms
17e. COSATI Field/Group
                    13/B
IB. Availability Statement
 ""limited
                                                  19.. Security Class (This
                                                    Repon)
                                                      UNCLASSIFIED
                                                  20. Security Class (This
                                                     UNCLASSIFIED
                                                   21. No. of Pages
                                                       115
                                                   22. Price
FOUM NTIC-tB 110-70)
                                                                     USCOMM-DC 40820-P7I V!

-------
              DISCLAIMER

This report was furnished to the Air Pollution
Control Office by  Monsanto Research Corporation
                   Dayton Laboratory
                   Dayton, Ohio  45407
in fulfillment of Contract No.  PH 22-68-12

-------
                                           MRC-DA-245
     APPLICABILITY OF CATALYTIC OXIDATION TO
      THE DEVELOPMENT OF NEW PROCESSES FOR
          REMOVING S02 FROM FLUE GASES
        Volume II - Experimental Program


            Contract No. PH 22-68-12



                MRC Job No. 6708



                   Prepared by

         R. E. Opferkuch, Project Leader
                   S. M. Mehta
                   M. G. Konicek
                   D. L. Zanders
          MONSANTO RESEARCH CORPORATION
                DAYTON LABORATORY
               Dayton, Ohio
                  January 1971
                  Submitted to

       Process Control Engineering Program
  National Air Pollution Control Administration
          Environmental Health Service
U.S. Department of Health, Education, and Welfare
              5710 Wooster Pike
            Cincinnati, Ohio 45277

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                             FOREWORD
The Intent of this volume is to present in an organized manner
the accumulation and assessment of all experimental and design work
accomplished during the execution of Contract PH 22-68-12 by
Monsanto Research Corporation, Dayton, Ohio.  In that this volume
represents the second phase of a three-phase program of concurrent
tenure, occasional reference to phases I and III may be seen at
various points through the text.  Further definition of these areas
may be found in the respective volumes, each of which is reported
under separate cover.

This final report is presented in three volumes in an effort to
make the material accessible on the assumption that it is of
practical value and therefore will be put to use.

Volume I is intended to contain all, and only, that material derived
from, or related to, the literature search.  Essentially all
information in Volume I is directly based on the literature.

Volume II presents data and information generated in the laboratory
and on the drawing board.

Volume III is an Indexed bibliography.

Finally, guidance through the three volumes is offered in the form
of the Foreword, Project General Summary and Tables of Contents in
each of the three volumes.

The authors wish to acknowledge the many helpful comments and
suggestions of the NAPCA Project Officer, Mr. George L. Huffman.
                                ii


                   • MONSANTO RESEARCH CORPORATION •

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                      PROJECT GENERAL SUMMARY
The salient features of the project are summarized briefly below.
Expansion and details are given in the texts of Volumes I and II.

The main objectives of the program were:

    a.  Search all available literature for pertinent information
        relative to the catalytic oxidation of sulfur dioxide, i.e.,
        active materials, mechanisms of catalysis, methods of
        application, equipment employed, etc.  Identify, describe
        and evaluate processes disclosed in the literature to have
        commercial potential for removal of sulfur dioxide from
        flue gas by oxidation.

    b.  Test, in the laboratory, candidate materials and methods
        suggested in the literature for potential application to
        removal of sulfur dioxide from flue gas by catalytic
        oxidation.

    c.  Identify at least one effective catalyst for the desired
        application and design a process for removal of sulfur
        dioxide from flue gas by catalytic oxidation and recovery
        of the sulfur value.

An intensive search of the literature revealed the following:

    a.  The transition metal oxides, notably vanadia, and platinum
        were the most commonly employed solid catalysts for practical
        conversion of sulfur dioxide to trioxide.   Nitrogen dioxide
        was the only practical gaseous catalyst noted.

    b.  Kinetic equations describing conversions of sulfur dioxide
        over vanadia or platinum catalysts were derived from data
        relative to commercial production of sulfuric acid, i.e.,
        high concentrations of sulfur dioxide.   There was nothing
        available to describe results at the comparatively low
        concentrations of sulfur dioxide found in flue  gas.

    c.  A number of processes were described as having  commercial
        potential for flue gas cleaning.  Comparative cost-
        performance evaluation of these oxidation processes
        eliminated all but two types as having realistic commercial
        potential, viz.,  one type based on vanadia catalyst and one
        based on nitrogen dioxide catalyst.

    d.  The most practical mode of recovery  of oxidized sulfur value
        from a use standpoint,  in this country, is production of
        fertilizer grade  sulfuric acid.
                   • MONSANTO RESEARCH CORPORATION •

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Laboratory tests and comparative,evaluation of commercial and
experimental catalysts indicated the following:

    a.  Commercial vanadia catalysts employed in production of
        sulfuric acid are effective in converting sulfur dioxide
        at concentrations found in flue gas.

    b.  Although platinum catalyst performance is essentially
        equivalent to that of vanadia, a cost-performance
        comparison indicated vanadia is ten times more effective.

    c.  No other candidate materials were shown to be as
        effective as vanadia, platinum or nitrogen dioxide in
        converting sulfur dioxide to trioxide,

    d.  Nitrogen dioxide was the only practical "low temperature"
        catalyst observed.

    e.  Using nitrogen dioxide a,s a catalyst, It is potentially
        practical to remove botjh sulfur dioxide and indigenous
        nitrogen oxides from flue gas simultaneously.

From preliminary process designs, and cost estimates, based on
laboratory data generated in this program, the following emerge:

    a.  Processes, based on vanadia catalyst, for oxidizing
        sulfur dioxide and removing it from power plant flue gas,
        as sulfuric acid, are likely to cost in the range of $12
        to $25 of capital per Installed Kw of power plant capacity.
        A large portion of the cost results from the need for
        corrosion resistant equipment.

    b.  Operating costs for vana,dia based propesses are likely
        to be in the range of P,5Q to 0.7*1 mills/Kw-Hr generated
        before sulfur value net back.

    c.  A substantial reduction in capital and operating costs
        are potentially available through a technique of sorbing
        the oxidized product gas from the main flue gas stream
        and recovering it separately.

    d.  The vanadia based processes are better suited to proposed
        new power plant installations than to existing plants
        because of numerous difficulties in retro-fit to existing
        power plants.
                                iv

                   • MONSANTO RESEARCH CORPORATION •

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                              VOLUME II

                          TABLE OF CONTENTS

                                                              Page

INTRODUCTION                                                    1

  I.   EXPERIMENTAL EQUIPMENT                                    3

      A.  Catalyst Test Unit                                    3
      B.  Temperature Control                                   3
      C.  Reactor                                               7
      D.  Analysis System                                       7

 II.   CATALYST PREPARATION                                     19

III.   CATALYST TESTING PROCEDURE                               23

 IV.   CATALYST TEST RESULTS AND DISCUSSION                     27

      A.  Commercial Vanadia Catalysts                         27
      B.  Commercial Platinum Catalyst                         33
      C.  Experimental Catalysts                               33
      D.  Molecular Sieve Adsorption Studies                   41
      E.  S02 Removal Characteristics of "Red Mud"             42

  V.   DEVELOPMENT OF A SIMULTANEOUS SOX-NOX REMOVAL PROCESS    45

      A.  Initial Laboratory Jnvestigations                •    45
      B.  Vapor Pressure Studies on the System                 48
      C.  Absorption Studies of NO and N02 into Sulfuric
           Acid              •                                  50
      D.  Proposed SOX-NOX (SONOX) Removal Process
           Description                                         54

          1.   General Description                              54
          2.   Trifunctional Absorption Tower           .        56

              a.   Bottom Section of Absorption Tower           56
              b.   Middle Section of Absorption Tower           56
              c.   Top Section of Absorption Tower              57

          3.   Acid Demister                                    57
          4.   Stripper                                         57
          5.   NO  Oxidizer                                      58
          6.   Acid-Acid Heat Exchanger                         58
          7.   Heat and Material Balances                        58

      E.   Comparison Between Tyco and Monsanto Modified
           Chamber Process                                     58
                  • MONSANTO-RESEARCH CORPORATION •

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                    TABLE OF CONTENTS (Cont'd)
         1.  General Description of Process Plow
             Differences                                      58
         2.  Detailed Process Comparison                      60
         3.  Detailed Comparison Using Identical Strippers    62

     F.  Remaining Questions to be Answered                   62

VI.  PROCESS DESIGN     '                                      65

     A.  General                                              65
     B.  Case I.  MRC/NAPCA Process - New 1400 Mw Plant       69
     C.  Case I-A.  Reversible Dry Absorption of SQ3          83
     D.  Case II - MRC/NAPCA Process - Existing 220 Mw
          Plant                                               92
     E.  Case III - Small Copper Smelter                     105

Appendix I - Commercial Catalyst Test Data                   115

Appendix II - Experimental Catalyst Test Data                129

Appendix III - Effect of Catalyst Particle Geometry on
               Pressure Drop in a Fixed Bed            >      135

Appendix IV - Reactor Design               .        .

Appendix V - Relationship of W/F Factors and
             Space Velocity

Appendix VI - Comparative Cost of Platinum and
              Vanadia  Catalysts                             153

Appendix VII - Sorptlon Isotherms for Molecular Sieves       l6l

Appendix VIII - Tyco Modified Chamber Process Report         171

Appendix IX - Vapor Pressure Apparatus and Procedures        183

Appendix X - S03-Product Acid Relationships                  187
                                vi


                  • MONSANTO RESEARCH CORPORATION •

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                          LIST OF FIGURES
                                                               Pag(

 1   Test Unit Schematic Plow System                             4
 2   Catalyst Test Unit Control Panel                            5
 3   Plow Controls                                               6
 4   Reactor Furnace Controls                                    8
 5   Glass Reactor and Preheater Section                         9
 6   Catalyst Bed and Temperature Sensing Assembly            • 10
 7   Reactor Furnace Enclosure                                 11
 8   Reaction Section Components                               12
 9   Furnace Sub-assembly                                      13
10   Reactor and Furnace Enclosure                             14
11   Process Stream Cooler Between Reactor and Chromatograph   16
12   Process Stream Analyzer                                   17
13   Analysis System Programmer                                18
14   Catalyst Pelletizer                                       20
15   Catalyst Pellet Surface (6000X) Before and After
     Heating at 800°F                                          21
16   Comparison of Experimental and Predicted Effects of
     W/F on Conversion at 800°F                                26
17   W/F Conversion Profiles for Catalyst 'A'                  28
18   W/F Conversion Profiles for Catalyst-E     .               29
19   Comparative Conversion Efficiency of Two Commercial
     Vanadia Catalysts at 800°F                                30
20   Comparative Conversion Efficiency of Two Commercial
     Vanadia Catalysts at 850°F                                31
21   Comparative Conversion Efficiency of Two Commercial
     Vanadia Catalysts at 900°F                                32
22   W/F Conversion Profiles for Commercial Platinum Catalyst  34
23   Comparative Conversion Efficiency of Platinum and
     Vanadia Catalysts at 900°F                                35
24   Effect of Time on Conversion Efficiency of Catalyst-B
     at 800°F, Constant Reactor Conditions (Run No. 6)         36
25   Effect of Time on Conversion Efficiency of Catalyst-B
     at 800°F, Constant Reactor Conditions (Run No. 7)         37
26   Conversion Efficiency versus Time Profile @ Constant
     Operating Conditions                                      43
                                 vii

                   • MONSANTO RESEARCH CORPORATION •

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                      LIST OF FIGURES (Cont'd)
                                                              Page
27   Laboratory Arrangement for NOX Scrubber                   46
28   Experimental Apparatus for Vapor Pressure Studies         49
29   Adjusted Vapor Pressure of NO and N02 Over Nitrose        51
30   Absorption Studies Apparatus                              52
31   Process Design Schematic                                  55
32   W/F Conversion Profiles for Catalyst 'A'                  68
33   MRC/NAPCA Process, Large New Power Plant Facility,
     1400 Mw, Process Plow Diagram                             70
34   MRC/NAPCA Process, Large New Power Plant Facility,
     1400 Mw, Plot Plan                                        71
35   MRC/NAPCA Process, Large New Power Plant Facility,
     1400 Mw, Elevation Drawing                                72
36   MRC/NAPCA Process, Large New Power Plant Facility,
     1400 Mw, Schematic Plan                                   76
37   MRC/NAPCA Process, Large New Power Plant Facility,
     1400 Mw, Instrument Flowsheet                             77
38   Effect of Product Credit on Operating Cost of MRC/-
     NAPCA Process (Large New Power Plant Facility)            84
39   Reversible Dry Absorbent Process, Process Flow Diagram    90
40   Effect of Product Credit on Operating Cost of
     Reversible Dry Absorbent Process                          97
4l   MRC/NAPCA Process, Small Existing Power Plant Facility,
     220 Mw, Process Flow Diagram                       .       98
42   Effect of Product Credit on Operating Cost of MRC/-
     NAPCA Process (Small Existing Power Plant Facility)      104
43   MRC/NAPCA Process, Smelter Facility                      107
44   Effect of Product Credit on Operating Cost of MRC/-
     NAPCA Process (Smelter Facility)                         113
45   W/F-Conversion Profile for Catalyst  'A' at 900°F         H6
46   W/F-Conversion Profile for Catalyst  'A1 at 850°F         117
47   W/F-Converslon Profile for Catalyst  'A' at 800°F         118
48   Effect of W/F on Conversion with Catalyst-B at 900°F     120
49   Effect of W/F on Conversion with Catalyst-B at 850°F     121
50   Effect of W/F on Conversion with Catalyst-B at 800°F     122
51   Effect of W/F on Conversion with Catalyst-E at 900°F     124
                                viii

                   • MONSANTO RESEARCH CORPORATION •

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                      LIST OF FIGURES (Cont'd)
                                                              Page
52   Effect of W/F on Conversion with Catalyst-E at 850°F     125
53   Effect of W/F on Conversion with Catalyst-E at 800°F     126
54   Effect of Catalyst Particle Geometry on Pressure Drop
     in a Fixed Bed                                           139
55   Sorption Isotherms for Molecular Sieve,  SK-20           163
56   Sorption Isotherms for Molecular Sieve,  SK-400          16^
57   Sorption Isotherms for Molecular Sieve,  SK-400          165
58   Sorption Isotherms for Molecular Sieve,  SK-410          166
59   Sorption Isotherms for Molecular Sieve,  SK-410          167
60   Sorption Isotherms for Molecular Sieve,  13X             168
61   Sorption Isotherms for Molecular Sieve,  SK-110          16.9
62   Baseline Process                                         175
63   Isothermal Scrubber                                      176
64   Catalytic Stripper                                       179
65   Catalytic Chamber Process                                180
66   Vapor Pressure Apparatus and Procedures                  184
67   Effect of S03 Concentration in Flue Gas on
     Dew Point of the Acid                                    188
68   Relationship of Flue Gas Dew Point to Acid
     Concentration in Condensate                              189
                                 ix

                   • MONSANTO RESEARCH CORPORATION •

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                           LIST OF TABLES
                                                              Page
 1   Catalyst Bulk Density                                     24
 2   Catalyst Test Data                                        25
 3   Summary of Economics for Vanadia  and Platinum Catalysts  38
 4   Effect of N02 on Total Conversion of Sulfur Dioxide       47
 5   Summary of Laboratory N203 Absorption Studies             53
 6   Heat and Material Balance - High Temperature Stripper     59
 7   Comparison of Monsanto and Tyco Processes (As Designed)   6l
 8   S02 Removal from Flue Gas                                 73
 9   S02 Removal from Flue Gas                                 74
10   Catalytic•Converter                                       78
11   Acid Recovery and Mist Collection Tower                   79
12   Acid Cooler                                               80
13   Acid Pump                                                 8l
14   Induced Draft Fan                                         82
15   Capital Cost Estimate Summary                             85
16   Equipment Cost Estimate Summary                           86
17   Working Capital Estimate Summary                          87
18   Operating Cost Estimate Summary                           88
19   S02 Removal from Flue Gas                                 91
20   Capital Cost Estimate Summary                             93
21   Equipment Cost Estimate Summary                           94
22   Working Capital Estimate Summary                          95
23   Operating Cost Estimate Summary                           96
24   S02 Removal from Flue Gas                                 99
                   • MONSANTO RESEARCH CORPORATION •

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                       LIST OF TABLES (Cont'd)

                                                              Page

25   Capital Cost Estimate Summary                            100

26   Equipment Cost Estimate Summary                          101

2?   Working Capital Estimate Summary                         102

28   Operating Cost Estimate Summary                          103

29   S02 Removal from Flue Gas                                108

30   Capital Cost Estimate Summary                            109

31   Equipment Cost Estimate Summary                          110

32   Working Capital Estimate Summary                         111

33   Operating Cost Estimate Summary                          112

34   Catalyst - A Test Results                                119

35   Catalyst-B Test Results                                  123

36   Catalyst-E Test Results                                  127

37   Experimental Catalyst Test Data                          130

38   Effect of Catalyst Particle Geometry on Pressure
     Drop in a Fixed Bed                                      140

39   Optimum Diameter of a Reactor Bed                        1^8

40   Summary of Economics for Vanadia  and Platinum
     Catalysts                                                155
                                  xi

                   •  MONSANTO RESEARCH CORPORATION •

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                          VOLUME II SUMMARY

During the first phase of this contract, which covered the
Identification and evaluation of existing and potential methods of
applying catalysis to the oxidation and removal of S02 from power
plant stack gas, a large quantity of information and data was
accumulated and assessed.  This evaluation revealed the need for
laboratory verification of the published data on promising oxidation
systems.  While the literature noted a large number of Individual
materials, and combinations of materials, as being capable of
converting S02 to 803, the experimental conditions, conversion
efficiencies, and extrapolated economics cited for nearly all cases
were impractical with respect to the flue gas application.

Essentially all these studies were related to conversion with
vanadla-based catalyst in gas streams containing 5-12$ S02, typical
of that in contact with sulfuric acid plant operation.  Accordingly,
it became necessary to provide laboratory verification of these data
under simulated flue gas conditions.

Commercially available catalysts were tested over a broad range of
conditions, resulting in the conclusion that vanadia-based materials
were the only practical compositions presently available with the
capability to remove SOa from power plant flue gas under normal
operating conditions.

A large number of potentially promising experimental compositions
were tested, but none was found which exhibited properties superior
to that of commercial vanadia-based catalysts for the pertinent task.
During this experimental investigation enough promising data was
gathered to merit further research in the area of new catalysts
which would operate at a lower temperature than those presently
available.

The development of a process which will simultaneously remove both
S02 and NOx was implemented experimentally as a logical outgrowth
of the experimental program.  A detailed description of this process
has been included in the text (Section V).  As this process bears
some similarity to the Tyco removal process, a comparison between
the two is also shown.

Preliminary catalytic oxidation process designs were developed for
application to a large new power plant facility, a large existing
power plant facility, a small existing power plant facility, and a
new smelter facility.  These designs include estimates of investment
and operating costs.   Since a great deal of data was obtained during
the execution of the experimental and design phases of the program,
most of it has been relegated to an appendix section at the back of
the volume.
                                 xli


                   • MONSANTO RESEARCH CORPORATION •

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                       VOLUME II CONCLUSIONS

1.  Vanadia catalysts available for commercial production of sulfuric
    acid are effective in converting SO?, in flue gas to SO3.

2.  Platinum catalysts available for commercial hydrocarbon reforming
    are effective in converting S02 in flue gas to SO3.

3.  -Although platinum catalyst is essentially equivalent in
    performance to vanadla catalyst, it is an order of magnitude
    more expensive.

4.  The presence of NOX in the flue gas stream had no apparent
    effect on catalyst conversion efficiency.

5.  No catalyst compositions tested, other than platinum and
    vanadia based compositions, were deemed satisfactory because
    of their instability to either SO,, or S03<

6.  No other combination of catalyst-promoter tested was as effective
    as the vanadia-based compositions for oxidizing S02 to S03.

7.  Catalytic oxidation processes are better suited to new plant
    design because of high costs associated with retro-fit into
    existing power station designs.

8.  Fleeting evidence in support of a proposed mechanism in the
    literature for vanadia catalysis, was obtained experimentally.

9.  No low temperature catalyst,  other than nitrogen dioxide,  was
    found'during the1 experimental search - however,  enough evidence
    was gathered to  merit  further investigation into low temperature,
    solid catalysts.
                                xiii


                   • MONSANTO RESEARCH CORPORATION •

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 INTRODUCTION

 The  theories  and principles  of  catalysts  are discussed  in Volume  I
 of this report.  The most  significant  conclusion emerging from  the
 discussions in Volume  I was that heterogeneous  catalysts by vandaia
 is the most effective  means  of  converting S02  to S03.   This  state-
 ment, however, does not specify a  formulation  of vanadia mixed. with
 anything else and does not suggest, in itself, how the  vanadia  is con-
 tacted with the sulfur dioxide.  Furthermore, there were, in the  litera-
 ture, several strong Indications,  or leads, to potentially new  cata-
 lysts that might operate at  lower  temperatures, or' that might be  more
 active per unit, weight.                        . .

 The  experimental effort, then, had two major objectives: (1) to seek
 "low" temperature or more active catalysts and (2) to define at least
 one  catalyst  around which a  process could be designed for removing
 S02  from power plant stack gas.

 In the following discussion, it is helpful to understand .the general
 terminology and structure of solid catalysts.  The basic parts  of
 a solid catalyst consist of:

      1.  Catalyst - active  material,  usually noble
          metals or transition metal oxides.

      2.  Promoter - usually an inorganic salt or
          salt mixture that  enhances the  activity
          of the catalyst.

      3.  Support - an inert material  such as Fuller's
          earth or alumina used to provide bulk to the
          mixture and increase surface area.

The  relative proportions of  the three  materials are by no means
 fixed but fall, roughly, into the following ranges, by weight:
                   Catalyst - 
-------
The ingredients can be combined by:

      1.  Dry mixing, pelletizing, calcining.

      2.  Adding enough liquid to make a dough and extrud-
          ing pellets, drying, calcining.

      3.  Dry mixing catalyst and promoter, calcining,
          grinding the "clinker", dry mixing with support,
          pelletizing.

      4.  Impregnating support with a solution of catalyst
          and promoter, drying, pelletizing, calcining.

It may not always be required to calcine.  When it is necessary,
however, the purpose is usually to convert the catalyst-promoter
mixture to the optimally active composition.  For example, the
desired composition of the promoter in a vanadia catalyst is the
pyrosulfate, whereas the promoter may initially be added as the
sulfate.

The number of permutations and combinations of catalyst pretreat-
ment operations promotes the impression that the field of  catalysis
is a "black art".
                    • MONSANTO RESEARCH CORPORATION •

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I.  EXPERIMENTAL EQUIPMENT

A.  Catalyst Test Unit

The system that was designed and fabricated for the purpose of test-
ing catalysts consists of the following functional sections':

          Preparation and metering of simulated flue gas

          Temperature control

          Reactor

          Analysis

Figure 1 is a schematic flow diagram of .the test stand.  Figure 2
is a view of the control panel.

The gas preparation section is designed to permit mixing and feeding
of simulated flue gas from cylinders, including calibrating gas mix-
tures.  Precise flow control is necessary here, and is provided through
multistage pressure regulation, precise flowmeters, and very sensitive
flow controllers.

The sample gas flow-control subassemblies, as.shown in Figure 3» con-
sist of two-stage regulation at the cylinder source followed by
rotameters with needle valve control.  Flow controllers are installed
downstream of the rotameters and are designed to control from con-
stant upstream pressure sources through variable.downstream-pressure
sources.  The flow rate set point is manually set by adjusting an
external valve handle.  This valve acts as a variable orifice to
provide a pressure drop across the high and low pressure chambers
of the diaphragm.  If the pressure .drop changes due to a change
in supply or outlet condition, the diaphragm will change to the operat-
ing point of an internal throttling valve to re-establish the pressure
drop across the manual valve.  The flow control system has the
capability of metering individual component flows, as well as multi-
component flows.  In the case of individual component flow, the
components are mixed downstream of the flow metering assembly through
a specially designed mixing chamber.

B.  Temperature Control

The mixed gas stream next passes through .a section of preheated line
prior to entrance into the reactor section, Tn the reactor section,
the gas is heated even further in an eight foot coil within the
                    • MONSANTO RESEARCH CORPORATION •

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                            Figure  1.  Test Unit Schematic  Plow .System
o
z
a
m
v>
m
O
I
O
o
3
TJ
O
O
Z
                                      L     oo<5<5
                      066
            A,  Gas Bottles
            B.  Flow Control Valves
            C.  Flow Meters
            D.  Pressure Gauges
            E.  Flow Controllers
F. Mixer
G. Inlet Gas Sample
H. Reactor Bypass
I. Preheat Section
J. Furnace Preheat Section
K.  Catalyst Bed
L.  Process Stream Cooler
M.  Condensables Trap
N.  Process Stream Vent
0.  Process Stream Analyzer
P.  Exhaust Stack

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Figure 2.  View  of Test Unit Control Panel
       • MONSANTO RESEARCH CORPORATION •

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•
        i


.
   „
«   . •
   •
   I.
     '
-
      .
               [
               I
                    .
                                   •
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             p
                               .
                               .
                        .

             .
                :

      Figure 3.  Plow Controls
  • MONSANTO RESEARCH CORPORATION

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reactor furnace enclosure.  Thermal input is controlled by two 220-
volt, 3-phase Variacs, one 3-phase contactor and a West temperature
controller, allowing exact adjustment of input power and temperature
control within +1°F.  Figure 4 shows the temperature control assembly.

C.  Reactor

Gas from the flow control system enters the reactor, Figure 5, through
an eight-foot spirally-wound,heating section ahead of the reaction tube
The hot gas then flows through the catalyst bed and out the exit•arm.
Figure 6 shows an enlarged schematic of the catalyst bed, thermo-
couples and supporting assembly,all of which may be positioned in,
or removed from, the reactor as one unit, allowing practical inter-
change of catalyst beds.  Figure 7 details the reactor-assembly
furnace-enclosure.  The enclosure is mounted on a 3/4-inch, 100-lb
steel plate which serves as a basic mounting surface as well as pro-
viding stability through a low center of gravity.  A high-temperature
blower extends through the base plate and into a stainless steel tank
12 inches in diameter and 24 inches in length.  Within this tank,
twelve 500-watt, high-temperature, strip heaters form an annulus sur-
rounding the glass reactor tube and provide .an 18-inch heated zone
along the length of the reactor and preheat assemblies.  Enclosed
in a squirrel cage housing, the fan blade exhausts up the outside
of the heat bank, and intakes along the center axis, providing a
uniform thermal environment for the catalyst bed.  The primary
tank enclosure is surrounded by a minimum of three inches of
FIBERFRAX ceramic insulation contained by a 30-gallon steel drum
mounted on the base plate.  Components of the reactor section are
shown in Figures 8, 9 and 10.

D0  Analysis System

The analytical system consists of an automatic, chromatographic pro-
cess stream analyzer.  The analyzing system has the capability to
resolve accurately S02, 02, and C02, with a multistage range
selector for S02, to encompass 0-0,2%, 0-0.4%, and Q-5% of this
component.   Other ranges are easily incorporated with the addition
of modular, attenuator cards in the system programmer0

The complete analyzing system consists of four major units: sample
conditioner, analyzer, programmer, and read-out device consisting
of an L&M series H recorder,,

The sample conditioner receives the raw sample from the process
stream and prepares it for introduction into the analyzer.  The
chromatographic analyzer separates the sample into its individual
constituents and, as they elute,  transmits signals to the pro-
grammer.   Signals are proportional to the concentrations of the
various components.  The analyzer utilizes a thermal .conductivity
detector with four filaments.
                    • MONSANTO RESEARCH CORPORATION •

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Figure 4.  Reactor  Furnace Controls
                  8
   • MONSANTO RESEARCH CORPORATION •

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                    BASKET a T.C. ASSEMBLY
                          \
SAMPLE GAS
   OUT
     PREHEATER
      SECTION
   SAMPLE GAS
        IN
CATALYST
  BED
 figure  5.   Glass Reactor  and Preheater Section
            • MONSANTO RESEARCH CORPORATION •

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          REACTOR TUBE
              CAP
          EXIT  GAS
         TEMPERATURE
               CATALYST
                BASKET
(25 mm O.D. x  40  mm hgt)

                SCREEN '
 INLET GAS
 TEMPERATURE
 BASKET
UPPORTING
  ROD
 CATALYST
    BED
(7/8 In. dia. x  1.5  in.  hgt)
     Figure 6.  Catalyst  Bed and Temperature Sensing Assembly
                                 10
                   • MONSANTO RESEARCH CORPORATION •

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INSULATION
CONTMNER
  HEATER
   BANK
       REACTOR
    ENTRANCE COLLAR
    TANK
  SUPPORTS
         HEAT
        MJMDf
    HEATER WIRE
      FEED THRU
     CONTROLLED
    ENVIRONMENT
      CHAMBER
                                                                   MOUNTING
                                                                    PLATE
HIGH TEMPERATURE
 FURNACE MOTOR
          Figure  7-   Reactor Furnace  Enclosure
                                 11
               • MONSANTO RESEARCH CORPORATION  •

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       •







         *
*•

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i
           1
     *r'
                '
                •
                •
        •
        .

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                     '•-!''•£&   t
               Figure  8.   Reaction Section Components

                                  12
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                                                •
                                            .
                                              •
                                            .
                                             •

                                                    •

f
              Figure 9.  Furnace  Sub-assembly
                             13
              • MONSANTO RESEARCH CORPORATION •


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Figure 10.  Reactor and Furnace Enclosure
      • WONSANTO RESEARCH CORPORATION •

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The programmer provides the control signals required to operate
the analyzer, and converts the analyzer-output signal into a
form compatible with the read-out device.  The sampling pro-
grams offered are:

       • Reactor inlet/outlet, sequential, iterative

       • Iterative inlet only

       * Iterative outlet only

       * Manual select

Various components of the analytical system are shown in Figure.s
11 through 13.
                                15


                   • MONSANTO RESEARCH CORPORATION •

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  Figure 11.  Process Stream Cooler
              Between Reactor and Chromatograph

       x
/*?
                       16
          • MONSANTO RESEARCH CORPORATION •

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Figure 12.  Process  Stream Analyzer
                 17
   • MONSANTO RESEARCH CORPORATION •

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               Figure 13.  Analysis System Programmer
KEPRODUCIBU
                                 18
                   • MONSANTO RESEARCH CORPORATION •

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II.  CATALYST PREPARATION

Laboratory equipment obtained for catalyst preparation Included
a  large capacity oven  (0-300°F) for drying samples, a small
high-temperature furnace (0-l830°F) for calcining operations;
and a pelletizing press  (Manesty), Figure 14.

The primary method of  sample preparation employed in catalyst
studies consisted of the following:

       • Weigh out constituents in fine powder form
         (pregrind if necessary)

       • Homogeneously dry mix constituents

       * Bake mixture overnight to remove free water

       ' Pelletize

       • Calcine

The support material most commonly used was calcium montmorillonite,
or Fuller's earth.  It was checked for its capacity to retain
molten alkali salts (promoters) by preparing a blend of it with,
for example, potassium blsulfate (20% by weight), crushed (to
approximately 300 mesh.  This blend was then pelletized arid heated
in a furnace 120°F above the normal melting point of the KHSOi,.
Subsequently, the pellets were removed and examined under magnifi-
cation, for signs of running or sweating of the KHS04.

In Figure 15, obtained with .the aid of the scanning electron micro-
scope, the upper photograph (6000X) shows the surface of a catalyst
pellet before calcining.  The large particles,   roughly 3 microns
in size, are, we believe, the mixture of catalyst and promoters.
The smaller particles can clearly be identified as diatomaceous
earth, the support.  The surface exposed to gaseous reactants is
highly Irregular and the pores exist as the space between  or
within the particles.

The lower photograph (6000X) in Figure 15 shows the surface of a
pellet after calcining at 800°F.  The large particles are absent
and the support material appears to have been rather uniformly
coated, presumably by the molten promoters in which the catalyst
is dispersed.

In some preparations, alumina was employed as support in a manner
similar to that for Fuller's earth.
                                19


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Figure 14.   Catalyst Pelletizer
               20
 • MONSANTO RESEARCH CORPORATION •

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Figure 15.
Catalyst Pellet Surface
(6000X) Before (Upper)
and After (Lower) Heating
at 800°F
                  21
      MONSANTO RESEARCH CORPORATION •

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A second modification of preparation consisted of mixing catalyst
and promoter, calcining, grinding the "clinker" and dry mixing
with the support and pelletizing.

A third modification consisted of soaking the support in a solu-
tion, or suspension, of catalyst and promoter, and evaporating
the solvent  (water) to deposit the salts and oxide on the sup-
port.
                               22


                  • MONSANTO RESEARCH CORPORATION •

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III.  CATALYST TESTING PROCEDURE

A complete test of a given catalyst would Involve determining the
conversion obtained by varying each of the following:

                         Temperature
                         Concentration of SOz

                         Concentration of Oz

                         Contact Time

                         Concentration of other constituents>
                         active or inert

The rate of reaction is also dependent upon the catalyst1 surface: area,
pore volume, and size as well as the variables listed above.

Testing of the effects and interactions of all the variables on
conversion for a large number of catalysts would have required an
inordinate amount of time and expenditure for the amount and
usefulness of the information obtained.  It was therefore decided
to pursue a more rapid but less comprehensive test program for
preliminary testing of catalysts.

Since the variation in SOz and Oz concentrations in stack gas is
not large and both are at low absolute concentration, average values
of their concentrations were employed and kept relatively constant
throughout the tests.  Thus, a feed stream of 0.3% SOz and 3$ Oz
would be representative.

Next, a standard set of test conditions was established as follows:

        Flow Rate - J»0 cc/sec (space velocity 2000-2500 hr"1)

        "Standard" stack gas composition*

        Initial run temperature - 600°F

It was then necessary to select a common basis for evaluating
different catalyst samples.   A first approach considered "contact
time" as the common base.  It soon became evident that this was
inadequate because the variation in bulk density, Table 1, among
different catalysts, resulted in widely different amounts of catalyst
showing the same contact time.
*Stack Gas Composition                  y ,  *
                      Nitrogen .......... 82.15
                      Carbon Dioxide. ... 14 . 70
                      Oxygen ............ 2. 80
                      Sulfur Dioxide.... 0.30
                      Nitrogen Oxides. . . 0.05
                                       100.00
                                 23

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                              Table 1

                         CATALYST BULK DENSITY
   Catalyst
   Code

   A

   B

   C

   D-l

   D-2

   E

   F

   G
Type

Vanadia

Pt

Vanadia

Vanadia

Vanadia

Vanadia

Pt

Pt/Fe
Weight    Volume    Apparent      Density
 (g)       (cc)     	(g/cc)	'
146.10

200.00

200.00
240

245

195
0.609

0.816

1.03
Sample  Catalyst  In  Powder  Form

Sample  Catalyst  In  Powder  F6rm

142.83     260             0.549

 64.80     145             0.447

 47.09      75             0.628
The 'Second approach considered the W/F value, i.e., the weight of
catalyst per mole of S02 per second.  When results obtained using
contact time as a basis were converted to a W/F basis, as in Table
2, they were found to fit reasonably well with the computer-predicted
curves derived from rate equations in the literature as illustrated,
for example, in Figure 16.  This, then, provided a common basis for
comparing different catalysts, namely: comparative conversion effi-
ciency at the same temperature and W/F value.  Further, the W/F
value is related to space velocity (See Appendix y)  thus permitting
reasonably valid extrapolation from laboratory test data to process
design.

In order to change the W/F value for a given weight of catalyst,
it was only required to change the gas flow rate.  Later, conversion
isotherms are presented for some commercial catalysts at various
W/F values attained in this manner.

In summary, the procedure employed for the comparative evaluation of
experimental catalysts consisted of determining the conversion
efficiency (% S02 converted) at the same W/F value under the standard
test conditions enumerated above.  In a number of instances, cata-
lysts were studied more broadly, as with some of the commercial
catalysts discussed below.
                   • MONSANTO RESEARCH CORPORATION •

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           Table 2



     CATALYST TEST DATA
Weight of
Catalyst, W Flow, F
Catalyst (gms) (cc/sec)
1.
2.
3-
1.
5.
6.
7.
8.
9-
10.
11.
12.
13-
14.
15.
16.
17.
18.
19.
20.
21.
A
(pellets)
A
(pellets)
A
(pellets)
A
(pellets)
A
(pellets)
A
(crushed)
A
(crushed)
A
(crushed)
A
(crushed)
A
(crushed)
B
B
B
B
C
C
C
C
C
C
D-l
5.8
5.8
5.8
5.8
5-7
6.0
6.0
6.0
6.0
6.0
8.5
8.5
8.5
8.5
9.7
9-7
9-7
9.0
9.0
9.0
7,6
lo
60
60
60
60
40
40
60
60
60
60
kO
60
60
10
60
60
60
100
. 60
60
W/F
gm sec/cc
0.145
0.0966
0.0966
0.0966
0.0950
0.150
0.150
0.100
0.100
0.100
0.142
0.212
0.1*42
0.142
0.242
0.162
0.162
0.150
0.090
0.150
0.127
Temp.
°F
800
800
650
900
900
800
850
850
650
900
850
850
650
900
850
850
650
850
850
900
900
Percent
Conversion
38*
29*
3.5*
46*
25*
52*
68*
58*
3.5*
68*
50*
65*
15-11*
70*
62*
51*
8-3*
18*
23*
53*
57*
              25



• MONSANTO RESEARCH CORPORATION •

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                                 Mori & Maenen 03
                      — —•————  Goldman «t. al.


                      	  Mars & Maeucn «1  425°C (800°F|


                             •   Cotolyit A, cruihtd
                      1.0               2.0

                      W/F, gm-iac/mole SOjXlO"6
                                        3.0
Figure 16
Comparison  of Experimental and Predicted
Effects of  W/F on Conversion  at 800°F
                              26
               • MONSANTO RESEARCH CORPORATION  •

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IV.  CATALYST TEST RESULTS AND DISCUSSION

From statements found In the literature, there was reason to believe
that commercially available S02 oxidation catalysts, designed for
use in sulfuric acid plants, would be effective with flue gas.  To
support this belief, commercial vanadia catalyst samples were ob-
tained from major U. S. producers.  One platinum catalyst, commonly
employed in hydrocarbon reforming, was also obtained.

A.  Commercial Vanadia Catalysts

In the particular case of vanadia catalysts, conversion efficiency
is rather poor below 790°F and above 930°F for a single pass con-
verter system due to greatly reduced activity at the lower tempera-
ture and the rapid shift in equilibrium at the higher temperature.
Consequently, three temperature levels were selected as representa-
tive of vanadium catalyst activity, namely:  800°, 850° and 900°F.
By experimentally defining  the W/F conversion profiles at these
temperature levels, a rather complete picture of the relative merits
of the various samples was obtained.  With the addition of catalyst
physical properties, it is possible to predict the weight of catalyst
needed for 90+# conversion of a given S02 concentration at constant
flow and temperature conditions.

Only two of the four commercial samples received were studied ex-
tensively.  A few data points were obtained with the remaining two
to support the assumption that they too would, in general, effect
conversion of S02 under flue gas conditions.

Approximately 20 experimental runs were required to determine the
W/F conversion profiles presented for the two vanadia catalysts.
The conversion isotherms are shown in Figures .17 and 18.  Examina-
tion of the curves indicates that 90$ conversion of S02 to S03 may
be attained at  typical stack gas concentrations (^0.3$) providing
the indicated temperature, flow, and catalyst charge requirements
are met (See  Appendix I for commercial catalyst  test  data).

As would be expected, there are point-to-point differences in the
performance of any two different commercial catalysts, shown in
Figures 19 through 21.  This is not only true "between manufacturers"
but also "within manufacturers".  Nevertheless,  the data support the
conclusion that the commercially available vanadia catalysts are
applicable, in a technical sense, to a process for removal of sulfur
dioxide from flue gas.  Furthermore, it is likely within the scope
of the manufacturers' present technology to adjust the performance
of these catalysts to give an incremental increase in efficiency
under flue gas conditions.  Process economics and competition will
probably motivate such adjustment.  Additional data relating to
vanadia catalysts can be found in Appendix III,  Volume I.
                                27
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                              CATALYST -A-
                              	    900°F
                                      850°F
                           	— •   800°F
                     1.0                 2D
                     W/F, gm.sec/mole SO2xlO"6
ao
Figure  17.   W/F-Conversion Profiles for Catalyst 'A'
                           28
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  100
Z
o
o
Z
iu
U
   30-
   20-
   10
                           CATALYST E
   900°F

   850°F


   800°F
                      1.0
 I

2.0
3.0
                      W/F, gm.tec/mole SO,
      Figure 18.   W/F-Conversion  Profiles  for Catalyst-E
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  100
   90-
   80-
   70-
Q
ui
>  60

O
u



8  30-
   40-
   30-
   20-
   10
        CATALYST A



	CATALYST E
                      1.0
         2.0
 i

3.0
                     W/F. Om.««c/mol« SOj
      Figure 19.  Comparative Conversion Efficiency of  Two

                   Commercial Vanadia Catalysts at 800°F
                             30
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  100
   90'
   80-
   70-
>  60-
O
u

8  50H
»-
z
iu
U
£  40-1
   30-
   20-
   10-
CATALYST  A
                             	CATALYST E
                      1.0
  I
  2.0
 i
3.0
                     W/F, gm-i«c/mol* SO,
      Figure  20.   Comparative Conversion Efficiency of  Two
                   Commercial Vanadia Catalysts  at 850°P
                              31
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100!
 10
                               CATALYST A


                       —	— CATALYST E
                   1.0
 i
2.0
 I
3.0
                  W/F, gm-i«c/mol« SOj
    Figure 21.  Comparative Conversion Efficiency  of Two
                Commercial Vanadia  Catalysts at  900°F
                          32
            • MONSANTO RESEARCH CORPORATION •

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B.  Commercial Platinum  Catalyst

Platinum catalyst designed for contact sulfuric acid plants is no
longer made commercially in this country.  The platinum catalyst
sample studied here was designed for hydrocarbon reforming.  How
much difference in composition or nature one might expect in
catalysts designed for these seemingly different purposes is not
known.  However, as indicated by the conversion isotherms in
Figure 22, this particular catalyst .gave a creditable performance.
Furthermore, a point-to-point comparison of the performance of
the platinum and vanadia catalysts at flue ;gas conditions presents
essentially the same pattern as that found in the literature for
the comparative performances of the two types -of catalyst at
"standard" contact plant conditions.  This can be seen in Figure
23 for comparative performance, at 900°F.

In contrast to the vanadia  catalyst, the platinum displayed an
unstable equilibrium at constant reaction conditions for a1con-
siderable period of time at. each set of conditions .tested.  The
effect of time on  the conversion efficiency is shown in Figures
24 and 25:  Sometimes as long as six hours was required before
a stable, reproducible conversion was obtained.

Although vanadia and platinum .catalysts would seem to be about
equivalent in performance, they are worlds .apart in cost.  A cost
analysis of these two catalysts, summarized in Table 3, indicates
that platinum must be at least ten times as effective as vanadia
to be comparable in cost with the vanadia catalyst.  Decline in
platinum catalyst cost is not likely in the near future in view
of the rising costs of labor and the metal .itself .(Details  in  Appendix  VI

C.  Experimental Catalysts.

Considering only the performance, characteristics of vanadia and
platinum catalysts, a major disadvantage with both, relative to
a flue gas process, is their high reaction temperature.  This
generates both economic and engineering difficulties in practical
application.  Consequently, it seemed logical to seek a catalyst
with an operating temperature nearer to thstt of the flue gas enter-
ing the stack, i.e., about 300°Fi  This temperature, however, is
well below the dew point of the sulfuric acid produced through
oxidation of the flue gas sulfur dioxide.  Thus,, a catalyst•system
operating at this temperature would soon be swamped with liquid
sulfuric acid.

A corollary to the "low" temperature approach is identification of
a catalyst which is more active throughout the present operating
temperature range than are commercial vanadia catalysts.  If a
                                33


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                                     —  900°F

                                     ....  850°F

                                     _  800°F
                i
               1.0
 I
2.0
 i
3.0
               W/F, gm.i«c/moU SOj
Figure  22.   W/F-Conversion Profiles for Commercial
             Platinum Catalyst
         • MONSANTO RESEARCH CORPORATION •

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lOOi
 10
                              	PLATINUM (CATALYST B)

                              	 VANADIUM (CATALYST A)
                    T
                   1.0
 T
2.0
3.0
                   W/F, gm.8«c/moU SO} *10~
   Figure  23.   Comparative Conversion Efficiency of
                Platinum and Vanadia Catalysts  at 900°F
                           35
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   40,
   36.
  32
  28-
at
in
> 24
Z
o
u


£201
t-
z
ui
U

S  16.
  12-
   4 .
                              CATALYST B
                      20               40               60

            ELAPSED TIME AT REACTION  CONDITIONS, minutei
      Figure
Effect  of Time on Conversion Efficiency of

Catalyst-B at 800°F,  Constant Reactor

Conditions (Run No. 6)
                            36
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Q
IU
Z
o
u
   65
  60
  55
  50-
  45-
O
to  401
at

a.
   30-
   25-
   20
   15
                            CATALYST B
                      1
                      30
I
60
 i
90
            ELAPSED TIME AT REACTION  CONDITIONS, minutes
     Figure 25.   Effect of Time on Conversion Efficiency
                  of Catalyst-B at 800°F,  Constant Reactor
                  Conditions  (Run No. 7)
                             37
              • MONSANTO RESEARCH CORPORATION •

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                           Table 3
     SUMMARY OF ECONOMICS FOR VANADIA  AND PLATINUM CATALYSTS^
                                V Catalyst         Ft Catalyst
Volume, Cu ft^1)                   30,300               3,020
Initial cost, dollars^2)        1,250,000          12,800,000
Regeneration cost, dollarsC2)   	           1,032,000
Capitalized cost, dollars(2)    3,920,000          20,380,000
NOTES:  (1)   Estimate of the volume of platinum catalyst that
              would be equal In cost to the estimated capitalized cost of
              vanadia  catalyst for 90$ conversion in flue gas.
        (2)   Estimates are based on the same W/F ratio for
              platinum and vanadia  catalysts.
        (3)   Assumptions used for comparison of platinum and
              vanadia  catalysts:
              General:
              a.  rate of return:  Q% per year
              Vanadia  Catalyst:
              b.  amount used for the oxidation of SOz in flue gas:
                  30,300 ft3
              c.  price:  $4l/ft3
              d.  packed density:  36.8'lb/ft3
              e.  catalyst life:  5 years
              Platinum Catalyst:
              f.  catalyst contains Q.5% platinum
              g.  packed density:  50 lb/ft3
              h.  catalyst cost:  cost of platinum + $2.75/lb
                  catalyst, (for manufacturing cost)
              i.  regeneration cost:  $0.75/lb catalyst
              j.  regeneration interval:  2 years
              k.  loss of platinum during regeneration: 2%
              1.  price of platinum: $110/oz (troy)
              m.  catalyst life:  30 years (including obsolescence
                  and/or abandonment)
                              38
                 • MONSANTO RESEARCH CORPORATION •

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catalyst were, say, 25-30$ more effective than vanadla throughout
the activity range, then at a temperature of 750°F, where vanadia
shows only 60-705& efficiency, the new catalyst may yield a desirable
90$ conversion efficiency.

Yet a third objective in testing experimental catalysts was to follow-
up suggestions, or clues, in the literature that could lead to new
catalytic compositions for efficient application to flue gas clean-
ing.

Observations made during the course of the commercial catalyst evalu-
ation studies appeared to shed some light on the mechanism of
vanadia  catalysis as proposed in the literature:
           S02   +   2V+5   +   O-2  *   S03   +  2V+"           (1)


           2V+1+  + 1/2 02 f 2V+5   +  O-2 (rate  con-      (2)
                                           trolling)


           V02,  V205  (solid)  	*V02, V205  (in  solu-  (3)
                                     tion in  molten alkali
                                     pyrosulfates)

In our catalyst test unit, at the end of a run, the mixed gas
stream is turned off and replaced by a stream of nitrogen at the
same bed temperature.  The nitrogen purges the system for roughly
an hour at temperature;  then the unit is shut down.  This opera-
tion, supplanting the mixed gas stream with nitrogen, has the ef-
fect of "freezing" the equilibrium mixture in Equation (2).  We
compared pellets of a catalyst "frozen", in this manner,  at 625°F
and at 800°F by placing them separately in distilled water.  The
pellets representative of the lower temperature were bluish in
color and produced  a blue solution, characteristic of V+1*.  Pel-
lets representative of the higher temperature were brown, charac-
teristic of V+5 and imparted no color to the water.  This qualita-
tive observation appeared to confirm the belief that at the lower
temperature, certainly the equilibrium mixture was predominantly
V+4*, but it raised the question of whether Equation (2) is inherently
temperature-dependent or whether it must act in molten solution,
which condition is temperature-dependent.  If it were solely the
latter, then a possibility of defining a system of low melting
promoters is offered.

There are salts which melt in the desired "low" temperature range,
and using some of these, fleeting evidence supporting the rationale
above was observed.  The short life of activity in the lower tempera-
ture region is attributed to instability of the salt either at
temperature or in the presence of S03.
                                39

                   • MONSANTO RESEARCH CORPORATION •

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Again we have seen vagrant activity of this type when, for instance,
residual chlorine in platinum catalyst is replaced with thiocyanate.
The increased activity of the platinum was short-lived because,it
Is proposed, of the reaction of sulfuric acid produced with the
thiocyanate to release thiocyanic acid.

Many catalyst compositions were prepared and tested in the search
for a stable low temperature, or more active high temperature
catalyst.  Included also were a number of experimental catalysts
from various vendors around the country.  These data are quite
extensive and, for that reason, are tabulated in Appendix II.  All
catalyst compositions beginning with the code letter (M) are
compositions prepared in our laboratory.  Other codes represent
experimental samples received from industrial catalyst vendors.
The latter were not accompanied by details of their compositions,
as this information is considered proprietary by the vendors.

It is interesting to note, throughout Appendix II, several com-
positions (designated by stars) which exhibit extremely low
temperature activity.  The apparent conversion efficiency reported,
while accurate, may be somewhat deceiving, as these data may not
represent optimum operating conditions of the catalyst..  However,
the data does indicate the possibility that these compounds have
potential as stolchiometric removal .systems similar to the limestone
injection process.  As mentioned previously, their performance may
be affected by several parameters not.yet defined.

It is felt, that the behavior nearly always observed, i.e., fleeting
low temperature activity, can be explained on the basis of the
product SOa reacting with the metal oxide catalyst forming sulfate,
or with the promoter, as with KHSOi|:


                 2 KHSO,, -I- S03  	»-   K2S207 + E2SOk


As the bisulfate is converted to pyrosulfate, the melting point of
the salt rises, with a corresponding decrease in activity at the
lower temperature of the test.

These observations are interesting in that they seem to support our
approach to a low temperature catalyst.   But, they also illustrate
the necessity of stability of the promoter system in the presence
of 863.   No material tested showed greater activity or stability
than vanadia under test conditions.

In the light of these cumulative observations, it is possible to
prescribe the requirements for a "low"  temperature vanadia catalyst,
which, in the end, resolve to the requirements for the promoter
salt.
                                40


                   • MONSANTO RESEARCH CORPORATION •

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 These  requirements  are:

    a.  Salt  (or salt system) should be completely molten
        and non-volatile in the  450°-500°F range.

    b.  Vanadia must .dissolve in molten salt and, probably,
        must,  ionize.

    c.  Salt  (or system) must be chemically stable in tempera-
        ture  range.

    d.  Salt  (or system) must be chemically stable in the
        presence of 02, S02 and S03.
These requirements assume that sorption and desorption of reactants
and products are not limiting, which is the current assumption for
commercial vanadia catalysts.

D.  Molecular Sieve Adsorption Studies

Molecular sieves had been shown in the literature to be useful in
many applications as heterogeneous catalysts and as vehicles in
studies of catalysts and catalyst mechanisms.  However, we could
find no publication which revealed specific studies to determine
their merit as support material for S02 oxidation catalysts.  As
the monodisperse nature of the zeolite pores and their extremely
high specific area were characteristics favorable for maximum
contact conditions between the S02 and active catalyst sites,
samples were obtained and used as support material for two previously
tested compositions.  No alteration of the catalyst characteristics
was seen.

Zeolite materials are also well known for their adsorption character-
istics with regard to the various constituents in stack gas, and we
wished to know if these characteristics would be an aid or
hinderance in their use as a support material.  Adsorption efficiency
profiles at four temperature levels between ambient and 250°F were
established for (4) different molecular sieve materials.  These
data are shown in  Appendix VII.    "       As regards materials
SK-410 and SK-400, both of which showed relatively high S02
adsorption efficiency, the effects of a 30$ reduction in adsorption
bed contact time are 'also shown.
                  •  MONSANTO RESEARCH CORPORATION •

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E.  SO? Removal Characteristics, of  "Red Mud"

During the course of our experimental catalyst testing program we
were requested by NAPCA to perform  evaluations of a material
commonly termed "red mud," which we were informed is a waste by-
product of aluminum ore processing  operations, for its potential
as a removal material for S02 in power plant stack gas.

Using the one sample sent to us, we ran a bed of the material
through our catalyst testing system and obtained conversion efficiency
vs. time profile as shown in Figure 26.   In a further attempt to
determine the constituents influencing the activity of red mud in
the presence of SC>2, a mass spectrometer analysis was performed on
a sample of exhaust gas from the reactor during the time period
shown in Figure 26.  No evidence of S03 was found.  However, these
results are inconclusive, as we were unable to determine the
sensitivity of the instrument to small concentrations of 803'.

Further tests were performed on the partially spent bed of red mud
which produced the decay profile illustrated in Figure 26.  Emission
analysis of the sample revealed the presence of the, following:

                          Fe — Low Major
                          Si — Low Major
                          Al — Low Major
                          Ca — Low Major (^10/S)
                          Na — 5 to 10%
                          Ti ~ 5 to 101
                          Zr — 0.3%
                          Cr — 0.1%
                          V  —• 0.07?
                          Mg — 0.03*
                          Mn — 0.03?

Subsequently, duplicate fusions were performed on the spent sample
to determine total sulfur content.  Results were 5-38 and 5-3^%
sulfur.  A water and acid insoluble residue resulted when the
fusion was taken up in water.  The best probability is that the
residue was a titanium containing material.

According to stepwise numerical integration of the absorption
versus time profile (Figure26) the sample of red mud absorbed 288 cc
S02 over a 95 minute contact time.  The density of SC>2 at metering
conditions is 2.6813 g/1.  Accordingly, the sample absorbed 0.772
gm S02.  Of this,  50% of the total would be sulfur, or 0.386 gms.
The sample weighed 9.82 gms, and of this an average 5-36/? was
determined as sulfur, or 0.526 gms.  Consequently, 0.140 gms of
sulfur were unaccounted for.
                  • MONSANTO RESEARCH CORPORATION •

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z
o
•z.
w

z
H
O

X
m
en
m
>  -t
a  LO
o
I

o
o
33
6
z
                   90
                   80
                g 70
                o 60
                LU

                OH
                O
                o
                   50
                o
                on 40
                    30
                    20
                                       Analysis Sample Taken (Mass Spec.)
                            ~K)     20     30     4&      50     60     70"

                                                    ELAPSED TIME, minutes
                                                                               80
90
100
                     Figure 26.  Conversion Efficiency versus Time Profile & Constant  Operating

                                 Conditions for "Red Mud"

-------
Duplicate fusions were performed on a fresh sample of red mud
with a result of Q.J2% sulfur, or 0.071 grts sulfur In the sample
Initially.  Thus, half of the missing sulfur was accounted for.
This left about 13% of the total sulfur unaccounted for.  However,
this Is within the accuracy of the overall technique.

These results Indicate that the activity of the red mud is due
to reaction between the S02 and various constituents of the sample,
with subsequent formation of stable inactive sulfur compounds
within the sample, causing the decrease in "apparent conversion"
noted in Figure 26.
                  • MONSANTO RESEARCH CORPORATION •

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V.  DEVELOPMENT OF A SIMULTANEOUS SOX-NOX REMOVAL PROCESS

It has long been recognized that high temperature combustion processes,
including electric power plants, produce oxides of nitrogen.  It
was therefore a logical and desirable extension of our experimental
program to check the effect various commercial catalysts had on
existing NOX concentrations in a power plant stack gas environment.
In addition, we felt that a process which would simultaneously
remove both SOX and NOX would have a decided advantage over a
process removing only one of the pollutants.

Examination of the well-known technology of the Chamber Sulfuric
Acid process indicated that nitric oxides could be scrubbed from
their carrier gas by 60° Be sulfuric acid - successful removal being
dependent upon the maintenance of equimolar concentrations of NO and
N02•  The resulting absorption of NO + N02 into sulfuric acid then
produces nitrosylsulfuric acid (NOHSOi,), this formation resulting
from the following reactions:

     • NO + N02 + 2H2S04 	>-  2NOHSOI, + H20                    (1)

     • 2N02 + S02 + H2S04 	>-  2NOHS04                         (2)

     • 2N02 + H2SOt» 	>-  NOHSOit + HN03                         (3)

The nitric acid which accompanies the formation of NOHSOi* In equation
(3) is then soluble in the large excess of sulfuric acid.

A.  Initial Laboratory Investigations

Realizing that a similar chemically favorable environment could be
produced.in the exhaust stream from our catalytic reactor, a small
glass scrubber was constructed and installed downstream of the reactor
as represented in Figure 27«  Sufficient N02 was then added to the
flue gas to meet the following minimum requirements:

     • 2 mols N02/mole unreacted S02 in flue gas

     • 1 mole N02/mole NO in flue gas

As we had no means available at this time to meter the amount of N02
into the system, an obvious excess was used (as evidenced by the
characteristic color of N02 in the combined gas streams).  The
rationale here was that the excess N02 would, react according to
equation (3)•

As soon as N02 was added to the reactor effluent,  white crystals
formed on the walls of the glass arm entering the bottom of the
scrubber.   The accumulation of these crystals shortly plugged the
                                45

                   • MONSANTO RESEARCH CORPORATION •

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               60°Be' I  H2S04
     ACID
 DISTRIBUTOR
 CONVERTER
  EFFLUENT
      CHAMBER
      CRYSTALS
     TO
CHROMATOGRAPH
                 ACID OUT
Figure 21.  Laboratory Arrangement for NOX Scrubber

          • MONSANTO RESEARCH CORPORATION •

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line.  These crystals were qualitatively  identified  as  "chamber
crystals" or the anhydride of nitrosylsulfuric  acid.  This was not
surprising, as the gas mixtures were dry.  More  interesting, however,
was the reduction of unreacted SC>2 in the  flue  gas.  At the time,
the catalytic converter was effecting about  90$  S02  conversion,
but the exhaust from the scrubber showed  95+% conversion.

The NC>2 feed point was moved to the scrubber and the experiment
repeated with no apparent formation of chamber  crystals.  The
exhaust from the scrubber was then analyzed  to  determine the effect
of NC>2 on the total conversion of sulfur  dioxide.  These data are
shown in Table  4.

                             Table 4

        EFFECT OF N02 ON TOTAL CONVERSION  OF SULFUR DIOXIDE


     Momentary total conversion without N02
         addition to scrubber                           89.1/8

     Momentary total conversion with N02
         addition to scrubber                     1)    9^.9$

                                                  2)  100.0$

                                                  3)    99.2$
                                                  4)    99.2$

                              v
Resulting product acid was shown qualitatively to contain nitro-
sylsulfuric acid.

The surprisingly high total conversion of  S02 resulting from the use
of N02 in the scrubber prompted further experiments.  The flue gas
was fed directly into the scrubber, by-passing the catalytic
converter.  Results of this experiment and a comparison with and
without the use of the converter are given below.

               Flow Scheme                    %  S02 Conversion

     Source -»• Converter -»• Scrubber                  89.1

     Source -»• Converter •*• Add N02 •*• Scrubber        99-5

     Source -»• Add N02 -»• Scrubber                    99.7
                                 47

                   • MONSANTO RESEARCH CORPORATION •

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B .   Vapor Pressure Studies on the System;  H2 SOU -H2 0-N2 0 3

Our success in obtaining "almost 1005? -SO 2~ oxidation by addition -of
N02 prompted the decision to design a process based on "modified"
Chamber technology.  Initial attempts to design such a process
pointed out the need for vapor pressure data on the system:
^SOit-^O-^Oa .  A literature search revealed two articles on the
subject - one by Berl (Ref. 1) in 1934 and another by Tseitlin and
Yavorskii (Ref. 2  ).  Berl's data covered the range of sulfuric
acid concentrations from 64.05? to 80.05?, at temperatures from 30°C
to 150°C and for N203 concentrations from 0 to about 1%.  The data
of Tseitlin and Yavorskii covered an acid concentration range from
81.2% to 84.14$ for corresponding N203 concentrations from 4.8% to
about 10.05? at 50°C, 70°C, and 90°C.

A brief vapor pressure study was also initiated to check the data
of these authors and to extend their work over the acid concentration
range of 64 to
A pictorial representation of the experimental apparatus is given in
Figure 28, which, incidentally, is quite similar to Berl's as given
in Appendix  9.  A 100-ml. round-bottom flask was filled approximately
half-way with the desired ^SO^-^Oa mixture.  The flask contained
a teflon coated magnetic stirring bar.  The flask was immersed in an
oil bath with temperature maintained, by an electric heater.  A (-10°
to 200°C) thermometer was suspended in the bath for temperature
measurement.  The flask was connected to one of two Hg manometers
by means of rubber and glass tubing - the choice of manometer
depending upon the expected vapor pressure.  A MacLeod gauge was
used to measure pressures from about 0.01 mm Hg to 5 mm Hg, and a
U-tube manometer for pressures between 3-700 mm Hg.  The system
was connected to a vacuum source equipped with a liquid nitrogen
trap and a dry ice trap which could pull a vacuum down to 0.007 mm
Hg.  There were two pinch clamps on the rubber tubing - one of which
was immediately above the round bottom flask, and a second on the
line to the vacuum source.

Our initial experimental results indicated that the data of these
authors was low by a factor of about 3«  However, we found that our
procedure used to determine the vapor pressure was faulty.  When this
fault was corrected, the data agree satisfactorily with the data of
Berl and Tseitlin and Yavorskii.

A description of the apparatus and procedure used to obtain the
vapor pressure data is given in Appendix 8 to point out the faulty
procedure and thereby prevent its recurrence in the future.  The data
is not included because it is of rather poor quality compared to
that of Berl, Tseitlin, and Yavorskii, and would only serve to
                                48

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o
z


z

o
m
to
PI
>
TO
O
I

o
o
3J
TJ
O
o
z
                      Heater & Controller,
                                                 Glass Tubing
100ml Flask
                                                         Stopcocks
                                                                Rubber Tubing & Pinch Clamps
                                                                               Hg
                                                                               Manometer
—To Vacuum&Liquid Na&Dry IceTrapsI
                                                            Magnetic Stirrer
                      Figure 28.   Experimental  Apparatus for  Vapor  Pressure Studies

-------
confuse the picture.  Instead, a summary of their data is given in
Figure 29-  Berl's data was plotted on a log P(NO and NOa) vs 1/T
plot for 1.0$ N203 mixture, and was limited to the following acid
concentrations:  64.0$, 67.0$, 68.5$, 70.0$, 73.3$, 75.65$, 78.0$,
and 80.0$.  The best line was drawn through each of the acid
concentrations.  As the lines appeared to intersect near  a single
point, the "best point" was chosen and all lines were drawn to
Intersect at that point.  The lines at 71 and 72$ were included
for convenience.  Lines above 80$ were obtained by extrapolating
Tseitlin's and Yavorskii's data back to 1$ ^03 and forcing the line
to pass through the intersection point.  The greatest deviation
from experimental data appears to occur at about 80$; however, these
deviations differed in direction for the two authors.  Berl's data
shows somewhat lower NO and NC>2 vapor, pressures (at 110°C and up)
than that shown on the plot, whereas Tseitlin's and Yavorskii's data
show somewhat higher NO and N02 pressures at 80$.

The vapor pressure data given in Figure 29 can be extended to other
N2©3 concentrations by multiplying by the ^03 concentration in
wt$.

The design calculations given in this report were based on the vapor
pressure data given in Figure 29.

C.  Absorption Studies of NO and NOg into Sulfuric Acid

A laboratory absorption column was constructed to test the kinetics
of absorption for NO and N0£ into H2SOi» and to compare experimental
absorption efficiencies with predicted efficiencies using vapor
pressure data.  Figure 30 is a schematic representation of the
column.  Two gas cylinders were used, one containing nitrogen and
the other NO, N02 and N2.  The amount of gas fed from each cylinder
was adjusted to provide NO and N02 vapor pressures of about 3 mm Hg
each.   These gases then passed into a mixing chamber to insure that
the gas fed to the column would be homogeneous.   The gas then
passed through an electrically heated tube and into the bottom of
the column.  The column and acid reservoir were heated by exterior
heating tapes, and insulated with 1/2 in. - 1 in. Pyrex wool.   A
description of the procedure used in this study is given in Appendix 9.

Table 5  is a summary of the data obtained from this study.  Eight
runs were made - the first three using a column packed with 1/2 in.
Raschig rings, and the last five with a column packed with cut Pyrex
tubing.  This tubing consisted of about equal portions of 5 mm and
6 mm tubing cut to 5 mm and 6 mm lengths respectively.

Examination of the data revealed that the efficiency of the 1/2 in.
Raschig rings to remove NO + NOz from the gas was about an order of
          less than the cut Pyrex tubing.
                                50


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1000
 0.1
    300 280 260 240  220  200   180    160    140
                                          120
                                                100
                                                       80
                                                                         40
                                  c (-fr SCALE)
     Figure 29-   Adjusted Vapor Pressure of NO and N02  Over Nitrose
                    • MONSANTO  RESEARCH CORPORATION •

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  Acid
  Storage
Acid Control Valve —
2.5' X 1.9'
   Column
Electric
 Heaters
Thermometer
 (

 Gas Sampler
FoVent
             Wet Test Meter
                        Thermometer
                            Gas
                              Sampler
                              LJ
               Cylinder
                        SCZZKI
                                  Mixing Chamber
                                                                Rotameters
                                                               Metering Valves
              N,-NO -NO,
              Cylinder
        Figure  30.   Absorption Studies  Apparatus
                     •  MONSANTO RESEARCH CORPORATION •

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                                               Table 5



                             SUMMARY OF LABORATORY N203 ABSORPTION STUDIES
Estimated
NO + N02


z
0
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H
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3
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en
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0
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TJ
0
TO
H
Press
Inlet
1.10

0.14

3.8
3.8

3.8

8.8

3.78

4.82
(mmHg)
Outlet
1.00

0.138

3.4
1.2

2.1

4.3

1.41

0.74
NO + N02
Removal
( %}
9-1

1.0

10.5
68.4

44.8

52.3

62.7

85.0
Approx.
Flow
Gas
cc/Min
12000

12000

12000
6000

6000

12000

12000

12000
Rate
Acid
cc/Min
15

80

100
80

33

35

55

60
Temperatures (°C)
Gas
In
119

120

126
26

26

82

25

110
Gas
Out
117

126

122
26

26

75

25

110
Acid Acid
In Out
129 75+

110

112
26 
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In an attempt to define  the  cause of the poor NOX absorption, a gas
sample of the N2-NO-N02  mixture was taken and submitted to I.E.
analysis.  The analysis  indicated that the sample had a respective
gas ratio of 2:1:0.4 as  N02/NO/N20.  This unexpected appearance of
N20 in low concentration was found to be a result of using a very
low grade NO supply.  The NO cylinder used to blend the mixture
contained about 25% N20  and  about 10$ N02.  Thus, of the total NO
and N02 present, only 50$ existed as a 1:1 ratio of NO to N02-

Since the removal of NO  + N02 from flue gas is optimum with a 1:1
ratio of NO to N02, and  falls off sharply as the ratio differs, it
is not surprising that the average removal of NO + N02 was about
63$ (avg. of last 5 runs).

Although the absorption  studies did not show the desired 99-6$ NOX
removal, they did point  out  the necessity of. maintaining a'lrl mole
ratio of NO to N02.  If  this ratio is not maintained closely, the
recovery of NO + N02 will be severely reduced.

D.  Proposed SOX-NOX (SONOX) Removal Process Description

1.  General Description

Utilizing the data.gathered during our laboratory investigation of
the dual purpose removal system, a process design was developed as
shown in Figure 31.   In  this process, flue gas enters the bottom of
a trifunctional absorption tower.  In the bottom section flue gas
is dewatered to ^0.2$ to 0.5$ H2.0 and cooled to 1?0°F by counter-
current contact with 160°F,  84$ H2SOi+.  The acid is then diluted
from 84$ to about 8l$ and heated to about 290°F.  The cooled flue
gas then enters the middle section of the tower where recycle. N02
is added.  The N02 oxidizes S02 to SOa and the resulting NO reacts
with residual N02 to form ^03.  The heat of reaction raises the gas
temperature about 20°F while the 803 and N203 are absorbed in, 80$
f^SOi* at 190°F.   The flue gas then enters the third section of the
tower where it contacts  the acid removed from the bottom section.
The gas becomes humidified and heated while the acid is reconcentrated
to 84$ and cooled to about l60+°F.  The acid may have to pass through
a cooler to bring the temperature back to l60°F, unless there is
evaporative cooling.

The acid removed from the middle section passes through a heat
exchanger and into a stripping tower.  Here the, dissolved ^03 is
stripped from the acid at 338°F with hot gas.   The denitrated acid
leaves the stripper, passes through the heat exchanger, a cooler,
and back to the middle section of the absorption tower.  A small
product stream is then removed beyond the heat exchanger.   The
stripper exit gas enters an oxidation chamber which provides
           time to effect 80+$ conversion of NO to N02.
                                54

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2
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H
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     U1
      25% Flue Gas to each of
          Four Towers
 Flue Gas
 300° F
7100 moles/min
 7.25%	*
 H20
                                                                         Clean Flue Gas
                                                                                270° F
                                                                                    50 moles/min
                                                                                   Product H SO
                                                                                                                        Product N02
                                                                                                                       -3.5 moles/min
                                                                                                         Recycle  N02
                                                                                                        -50 moles/min
                                                                                     -10% NO,
                                                                                   122°p
                                                                                                                                Cooling Water
                                                    80%* 900
                                                         moles
                                                       (N02)
                                                     (-12.5 moles/min)
                                                                                                                             40psig Steam
                                                                                                                                50moles/min
                                                                                                      3600 moles/min
                                                                               81% Acid-290° F
                                                                               lOOOmoles/min
                                                                                                      Condensate
                                                                                                      80 moles/min
         Vaporized
         Condensate
 Product Acid Out
~50moled/min
                                                                    ONE OF FOUR TOWERS
                                                                              ONE STRIPPER
                                                Figure  31.   Process  Design  Schematic

-------
Water vapor is condensed by cooling the gas to 122°P with water.
The NOa-rich gas stream, minus a small product stream, is then
recycled to the middle section of the tower.  The condensate from
the oxidizer is revaporized and fed back to the stripper.  In
addition, make-up water is added to the stripper as steam.

2.  Trifunctional Absorption Tower

    a.  Bottom Section of Absorption Tower
Vapor pressure studies on the system I^SO^-HaO-^Oa revealed the
desirability of operating the absorption section of the tower, at
as low a temperature as possible and with an acid concentration of
80+/K.  In order to meet this requirement, the flue gas must be cooled
to a suitable temperature (^170°F) and must be dewatered to the
equilibrium composition of H20 vapor over 80% sulfuric acid at ^190°?.
The requirement that the gas must be dry arises from the fact that
absorption of H20 into the acid would be accompanied by a decrease
in acid concentration and a rise in .-temperature from the heat of
absorption - both detrimental to NO + N02 absorption.

The following assumptions were made in designing this section of the
tower:

        1.36 pounds of 84$ HaSOit/pound of flue gas

        The temperature of the acid leaving the tower will be
        10°F cooler than the entering flue gas temperature

     b.  Middle Section of Absorption Tower

In this section the S02 in the flue gas is oxidized to SOa by NC>2
(obtained from the recycle stream), thereby permitting removal of
the sulfur as H2SOi+.  The amount of N02 recycled is controlled to
affect a 1:1 ratio of NO to N02 after all the S02 is oxidized.  The
temperature of the recycle N02 stream is about 122°P and the heat
of S02 reaction is sufficient to affect a 20°F rise in the flue gas
temperature.  Eighty (80) percent H2SOit is used to scrub the NO + N02
gases from the flue gas.  At these operating conditions, approximately
0.88 pounds of acid/pound of gas will remove 99.6!? of the NO + N02
with 15 theoretical trays.  The flue gas leaving this section
contains about 0.25% H20, 0.003? S02 , and about 0.005% NOX, and the
acid leaving this section contains about 1.08 wt# N203 (3-92$ HNS05).

Here, the following assumptions were made:

        The oxidation of S02 to S03 is instantaneous

        Fifteen theoretical trays are desired

        No oxidation of NO takes place in the absorber.
                                56

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     c.  Top Section of Absorption Tower

The function of this section is to utilize the dry clean flue gas
to concentrate sulfuric acid from Bl% to 84$.  As the Ql% acid
enters the top of the section  (from the bottom of the bottom section)
at 690°F, the gas passes countercurrent to the acid and removes
H2
-------
5.  NO Oxldizer

The function of the NO oxidizer is to provide sufficient residence
time to allow the oxidation of NO to take place.  Only one tower
(about 35' x 30' diameter) is required.  This provides 9 sec.
contact time.

Assumptions made in this unit:

        80+/? conversion of NO is desired

        The gas is cooled to 122°F

        The reaction constant is 1.5 x 101*  (l2/mole2- sec).

6.  Acid-Acid Heat Exchanger

The function of the acid-acid heat exchanger is obvious.  The size
of the heat exchanger (surface area) depends upon the heat transfer
coefficient and the AT.  For an assumed heat transfer coefficient of
150 BTU/hr'ft2-°F and a AT of 14°F, the required surface area is
3^5,000 ft2.

Assumptions made for acid-acid heat exchanger:

        Temperature driving force is 14°F

     •  Heat transfer coefficient is 15 BTU/hr-ft2•°P.

7.  Heat and Material Balances

The material balance for this process is given in Table 6 .  .The
flow rate of each compound in a given stream is shown in lb-moles/-
min.  The temperature and pressure of each  stream is included if
significant.

E.  Comparison Between Tyco and Monsanto Modified Chamber Process

1.  General Description of Process Flow Differences

The description of the Tyco modified chamber process was taken directly
from a report issued by Tyco representatives at the Second Annual
Control Process Contractors' Meeting (NAPCA) held on June 11-13, 1969,
in Cincinnati,  Ohio.   A copy of that report is included in Appendix VIII

The Tyco process basicly rests on two important design features:
(1) the absorption of SOa and N203 (NO + N02) is accomplished in an
absorption tower operating isothermally at  250°F using 80/J I^SOit,
',_) tne absorbed N20a (HNSOs) is oxidized in the liquid phase over
a charcoal catalyst and the resulting N02 is removed by the gas
sweeping through the column.
                                58

                   • MONSANTO RESEARCH CORPORATION •

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             Table 6



HEAT AND MATERIAL BALANCE - HIGH TEMPERATURE STRIPPER





*
z
O
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3>
•z.
H
O
a
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a vo
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0










Stream
Flue Gas
Gas Exit
BTM
Gas ABS
Exit

Clean
Flue
Recycle
N02
Prod. NO 2
Oxldlzer
Eff.
Stripper
Gas Out
Strip Air
ABS BTM
Acid Out
ABS Top
Acid Out
ABS Feed
Acid
ABS Acid
Out
Strip
Acid Out
Prod. Acid
Cond.
Oxldizer
H20
Make-Up
H20

N2
5318
5318

5578


5578

260

16.5
276.5

276.5

276.5
__

__

__

_

	

_ _

,„

—


03 C02
199 1044
199 1044

258 1044


258 ion

59

3-7 —
62.8

73-5 —

73-5 —
— —

— —

. 	 	

— —

— —

-- —

	 __

—


H20 S02 N02 NO H2SOU N203
515 21.3 0.35 3.20 —
21 21.3 0.35 3.20 —

21 0.3 0.17 0.17 —


515 0.3 0.17 0.17 —

28.9 — 45.45 5-05 —

1.8 — 2.88 0.32 —
30.7 -- 48.33 5-37 —

112.3 — 26.85 26.85 —

—
3106 — — — 2483

2612 — — — 2433

2066 — — — 1504 0.08

2075 — ~ — 1525 26.93

2094 -- — — 1525 0.08

28.8 — — — 21.0

81.6

51.6 — _

Temp.
(°F)
300
170

190


270

122

122
122

338

565
290

160V
100
203/
190
190/
323
338/
203
203

122/
380
380

Press.
(mm Hg) Comments
785
780

775


770

790

790
790 Remove 2.86xl06 Btu/min

800

•\-900
__

Cooler may be required

	

—

—

__

40 pslg
} ^.$800,000/yr
40 pslg


-------
Of the two features the second Is more important.  It makes it
possible to operate the stripper at the same temperature as the
absorber.  Thus, there is no necessity to heat the acid in order to
raise the NzOa vapor pressure and permit desorption, and later cool
it to the absorber temperature.  The rest of the process consists of
(1) recovering nitric and sulfuric acids, and (2) preparing the flue
gas to enter the absorption tower by cooling to 250°F and reacting
NOa with S02.

The Monsanto process differs from the Tyco process on both important
design features.  The absorption tower is trifunctional rather than
monofunctional.  In the first section of the tower the gas is cooled
to about 170°F and dried to approximately 0.2-0.5? HaO with 84$
sulfuric acid.  In the second section, N02 is added, oxidizing the
S02 to S03 and the 80s plus NO and N02 are absorbed with 80$ H2SOi+
at 190°F.  In the third section the flue gas is warmed and moistened
by passing through the hot, humid acid removed from the bottom of
the first section.

The stripper is also different.  Assuming that neither charcoal or
any other catalyst will effect the liquid phase oxidation of NO,
the recovery of NO + N02 is accomplished by raising the acid
temperature from 190°F to 338°F.  At this, temperature, the vapor.
pressure of NO + N02 is about 30 times higher than it was at 190°F,
the absorbing temperature.  Thus, the stripping of NO = NOa should
be accomplished with less than 5% of the flue gas volume.

2.  Detailed Process Comparison

Table 7 is a direct comparison between the Tyco and Monsanto,processes
The comparison is made as the two processes were described.  Tyco's
stripper is catalytic and Monsanto's stripper is high temperature.
The required absorption acid flow rates were calculated using the
vapor pressure data shown in Figure 29 as a common basis.  The major
differences between the two processes are the following:

        Tyco process requires about twice the acid flow rate in
        the absorber as Monsanto's process

        Monsanto process is circulating 1.08 times as much acid
        as Tyco process

        Monsanto absorption towers will probably cost 2 to 2.5
        times as much as Tyco towers

        Monsanto absorption tower is about 3 times higher than
        Tyco towers

        Monsanto process uses 24,000 GPM cooling water while Tyco
        uses 17,000 GPM.
                                60

                   • MONSANTO RESEARCH CORPORATION •

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                                                   Table 7

                           COMPARISON OF MONSANTO AND TyCO PROCESSES (AS DESIGNED)
      Subject
                                  Monsanto
      Gaa'Cooling
 9.


10.

11.



12.


13.


ID.
15.


16.
      Product Acid Concentra-
      tion

      Relative required add
      flow rate for NiOa
      absorption

      Stripper Operation
      Stack Oas Exit
      Temperature

      NO Oxidation Chamber
      Total Relative Acid
      Circulation Rate
      (Absorber Stripper)

      S02 Reactor
Acid Heat Exchangers
Required

Acid Coolers

Maximum HNSOs cone.
In absorber (at
minimum L/C)

Relative Absorber Tower
Construction Difficulty

Stripping Tower Con-
struction Difficulty

Catalyst Losses
Pressure Drop by Flue
Oas

Cooling Water Require-
ments
                            Uses a flue gas cooler to remove
                            1.3xl06 Btu/min as low-grade steam

                            Area of heat transfer • 90,000
                            ft2 (cools gas to ^278°P)
                            80S


                            1.0 (^1814,000 Ibs/mln)
                            Uses temperature increases to strip
                            N0+N02.  Air Is used as stripping
                            gas (possibly with some flue gas).
                            No catalyst Is assumed.  Required
                            gas rate Is 350 moles  (at least
                            120 must be air)
                                  270+°P
                            A tower 30* 0 by 36" (provides
                            •>-9 sec contact time)
                            1.08 (have three separate
                            absorption sections)
                            Done In absorber
One acld-to-acid exchanger with
327,000 ft2

8 (total 15,000 ft2)

6.08
2-2.5


1.0


None






•••26" HjO


24,000 GPM
                                                                           Tyco
                                         Cools all flue gas to 250°F;
                                         removes heat of SOj oxidation;
                                         removes heat of NO oxidation
                                         Heat removal =
                                            2.9xl06 Btu/min - cooling
                                            O.SxlO6 Btu/min - SOz ox.
                                            0.6xl06 Btu/min - NO ox.

                                         80S  (limited to 92* maximum)
                                         2.0 (•v.381,000 Ibs/mln)
                                         Uses charcoal as a catalyst to
                                         oxidize N203 to 2 N02.  Flue
                                         gas Is used to oxidize NO and
                                         strip N02 from acid.  Requires
                                         minimum of 850 moles/mln of
                                         flue gas to provlda sufficient
                                         02 for oxidation (if no air Is
                                         used)
                                         None, required If charcoal
                                         catalyzes NO oxidation in
                                         stripper.

                                         1.0 (one absorption section)
                                         Done prior to entrance to
                                         absorber In a tower (unknown
                                         dimensions)

                                         None
None

1.3*



1.0
There appears to be some
indication that NO] oxidizes
carbon to CO or C02.   Since this
would be accompanied by NO:
reduction, additional oxygen
Is required.

•v.20" H20
17,000
                                                      61
                                •  MONSANTO  RESEARCH CORPORATION  •

-------
        Monsanto process requires an acid-acid heat exchanger
        (3^5,000 ft2) and 8 acid-water heat exchangers (total
        of 45,000 ft ).   Tyco process requires no acid heat
        exchangers.

3.  Detailed Comparison Using Identical Strippers

If a catalyst does exist which will effect the oxidation of
to N02 in the stripper,  then the absorber is no longer required.
The process flow becomes much simpler.  There is no need to have
acid-acid heat exchange or acid coolers, and the absorption tower
can be operated as a monofunctional tower.  The dewatering and
humidifying sections can also be eliminated.

The Monsanto process and the Tyco.process thus become very similar.
The major differences are the following:

        Monsanto oxidizes 862 in the absorber
          Tyco oxidizes S02 in a separate chamber

        Tyco process operates at 250°P requiring flue gas cooling
          Monsanto process operates at . 300°F requiring flue gas
          cooling

        Tyco product acid concentration will be 80$
          Monsantofs will be 87%.

P.  Remaining Questions to be Answered

        Will evaporative cooling eliminate the need to cool the
        acid leaving the third section of the absorption tower?

        What is the maximum gas flow rate that can be tolerated
        in the absorption tower?

        Will NO oxidation take place in the stripper and/or
        absorber and eliminate the need for an oxidizing chamber?

        What is the minimum acid rate required to cool and dewater
        the flue gas?

        Is it more economical to increase the size of the acid-acid
        heat exchanger or to increase the temperature of the
        stripping air?

        Can flue gas be used as part of the stripping air (and
        H20 requirement)?

        Is it more economical to operate the stripper at high
        temperature and low air rate or low temperature and.high
        air rate?
                                62

                   • MONSANTO RESEARCH CORPORATION •

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What is the height of packing required in each of the
three sections of the absorption tower and stripping tower?

What will the pressure drop be through the absorption tower?

Is it necessary to monitor the NO/N02 ratio in the absorption
tower?  If so, how?

What materials of construction are required?

Is there a catalyst that can be used in the stripper that
will permit oxidation of N203 to N02 and thereby facilitate
stripping of NOX at lower stripping temperatures?
                        63

           • MONSANTO RESEARCH CORPORATION •

-------
                             REFERENCES
1.   fieri, Z. Anorg. Allgem. Chem. 202, 113-34  (1931).

2.   Tseitlin and Yavorskii, Journal of Applied Chemistry/U.S.S.R.
    39 (5) (1966).
                   •  MONSANTO RESEARCH CORPORATION •

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VI.  PROCESS DESIGN

 A.  General

As a result of laboratory testing of catalysts of various types, a
process design evolved based upon commercial vanadia catalysts.  It,
is necessary to note that the designs presented below were developed
solely from laboratory data supplemented with pertinent information
from the literature and various equipment vendors.  There was no
piloting of the processes under the contract.  It is also necessary
to note that the basic process.proposed here and labeled MRC/NAPCA
is superficially quite similar to published descriptions of Monsanto
Company's Cat-Ox** process.  The two were developed completely
independently of each other.  Published details of the Cat-Ox© process
only appeared toward the end of the present design effort.  Although
such information was too late to be influential in the designs shown
below, it did serve very conveniently as a means of checking the
reasonableness of these designs at a number of points.

Four process designs are presented.  The key step in each is
conversion of S02 to SC>3 over a supported vanadia catalyst at 850°F.
The basic process consists of such conversion followed by heat
recovery steps and then condensation of the sulfuric acid produced
for recovery.   A minor departure from the basic derives from the
available temperature of flue gas.  This normally is about 300°F.
for existing power boilers.   The catalytic converter, however,
requires feed gas at 850°F.   Consequently, the basic process applied
to an existing power plant requires adding provision for preheating
the gas fed to the converter.

In applying the process to a proposed power plant, provision can be
made in the power plant design to provide gas to the converter at
850° - 900°F.  without the additional heating step.

A major departure in design, shown below, from the basic process
consists in the mode of S03  recovery proposed by the Gallery Chemical
Company.

The four cases of process designs presented, then, are intended to
provide a view of-the basic  process, with various modifications,
under most probable applications.

        Case I.     MRC/NAPCA process applied to a large new
                   (1400 Mw) power plant (no separate preheat)

        Case I-A.   High temperature oxidation by solid catalysis
                   followed  by reversible dry absorption of S03
                   (Gallery  Chemical Company modification)
                                65


                   • MONSANTO RESEARCH CORPORATION •

-------
        Case II.   MRC/NAPCA process applied to a small (220 Mw)
                   existing power plant (separate preheat)

        Case III.  MRC/NAPCA process applied to the off-gas from
                   a reverberatory furnace of a small copper
                   smelter

Case I is illustrated in greater detail than the other three cases.
Capital and operating cost estimates are presented for all four
cases.

Cost estimates are based upon the following general assumptions for
all four cases.  Special assumptions pertinent to a particular
process are presented along with the process descriptions below.

        Flue gas and off-gas analysis:


                      Power Plant                   Smelter
Component.       Flue-Gas (% by Volume)        Off-Gas (% by Volume)

   N2                   74.9                          76.75

   C02                  14.7                           3.48

   H20                   7.25                          8.94

  '02                    2.8                           9.38

   S02                   0.3                           1.45

   NOX                   0.05

   Ply Ash               0.2 (by weight)


        Following ratio factors were used for estimating the fixed
        capital cost:

   Piping          50$ of purchased equipment cost (except for
                   large new power plant facility)

   Instruments     15$ of purchased equipment cost

   Insulation      12$ of purchased equipment cost

   Electrical      10$ of purchased equipment cost

   Building        10$ of purchased equipment cost
                                66

                   • MONSANTO RESEARCH CORPORATION •

-------
Land and yard
 improvements

Utilities

Engineering &
Construction

Contingency &
Contractor's
Pee
    of purchased equipment cost


    of purchased equipment cost

2055 of physical plant cost


20% of direct plant cost
     Catalyst Life

     Catalyst Cost
            5 years

            $l.l»5/liter
     Amount of vanadia catalyst is based on W/F - conversion
     profile for catalyst "A" at 850°F (Figure 32) plus 10$
     excess catalyst

     Size and shape of catalyst   3/8 in, cylinders
     Rate of Return

     Direct Labor

     Supervision



     Maintenance

     Plant Supplies

     Utilities
     Payroll Burden

     Plant Overhead
            o% per year

            $3.00/hr.

            $7800 and  $12,000 annually for first-
            line supervisors and area super-
            intendents, respectively

            5% of the  fixed capital investment

            15% of the maintenance cost

            a)  steam  — 50
-------
o
ui
H-
et
ut


Z

O
u
z
UJ

U
oc
IU
   30
   20
   10-
    CATALYST  -A-



	   900°F



            850°F



            800°F
                        1.0                  20


                        W/F, gm-sec/mole SO2><10'6
                                  ao
       Figure  32.   W/F-Converslon  Profiles  for Catalyst 'A1
                               68
                 • MONSANTO RESEARCH CORPORATION •

-------
      •  Depreciation             10%  of  the  fixed  capital  investment

      •  Taxes                    2% of the fixed capital investment

      •  Insurance                1% of the fixed capital investment


Major equipment costs  for  a  large new  power plant  facility  (MRC/NAPCA
process) were  substantiated  by  quotations  from equipment vendors.
For other  cases, equipment costs were  estimated  by using cost data
gathered by Monsanto's Central  Engineering Department and cost data
from authoritative publications.  Piping cost for  the large new power
plant facility was estimated from labor and material take-off.  For
this type  of estimation, process flow  sheet (Figure 33) and plan
and elevation  drawings (Figures 31*, 35) were used.

Cost of equipment or other facilities  normally required for usual
power plant operation  is not incorporated  in the capital investment
or operating cost estimates.  Only  the additional  and/or incremental
costs for  air  pollution control are estimated.

B.  Case I.  MRC/NAPCA Process  - New 1*100  Mw Plant

For a new  power plant, the design assumes  the availability of flue
gas at  a temperature of 850°  -  900°F.  A flow diagram for one train
of this process is shown in  Figure  33-  There are  four trains for
the 1*100 Mw station.  Table  8 presents an  overall material balance
and Table  9 shows the balance for a single train.

Hot gas from the boiler is relieved of 99+% of its dust burden in
a high  temperature, electrostatic precipitator.  The near perfect
efficiency of  fly ash removal is necessary to minimize clogging, or
blinding,  of the catalyst bed in the next  step.  The gas then flows
through fixed beds of vanadia   catalyst where 90% of the sulfur
dioxide is oxidized to sulfur trioxide.  The design of the
catalytic  converter required  optimizing pressure drop, through the
catalyst bed, with tower diameter.   Tower  diameter establishes gas
velocity which affects pressure drop.  The latter relates directly
to operating cost.   The converter design evolved consists of a set
of shallow catalyst beds in  parallel.  Gas flows up through the
beds, as indicated in Figure  35, to facilitate periodic cleaning
beds to remove fly ash.  There are four such converters,  each
with 13 beds for the gas handling capacity at this station size.
The program employed for converter optimization is given in Appendix
IVto this volume.   Further,  in Appendix III are the program and
results establishing the relationship between catalyst particle
geometry and pressure drop through the catalyst bed.

Gas leaving the converter is then cooled to about 500°F.   In a new
power plant cooling is accomplished by stepwise passage through an
                                69

                   • MONSANTO RESEARCH CORPORATION •

-------
M



O
n
(A
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•a
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z
                                                                                                  f'f. 1
         CATALYTIC
         CONVERTER
 FROM
ELECTROSTATIC
PRECIPITATOR
                                                                                                   200-225°F
                                                  TO
                                                 BOILER
                                                                 COOLING WATER
                                                                 WATER TO ECONOMIZER
                                                          FURNACE
                                                     PUMP   ACID
                                                           TO
                                                          STORAGE
         K-i
                                                                    . ^-1	L.__
                                                                         I
                                                                      -- f-
                                       *H-B

                                      , ..'-.ir

                                      OK«»--  I
                                     H
                                                                                    MRC-fMPCA PROCESS
                                                                                  LARGE NEW POWER PLANT FACIL/Ty
                                                                                          1400 MW
                                                                                  SCAM
                                                                              PROCESS FLOW DIAGRAM
                                                                                                                    OF
        Figure  33.   MRC-NAPCA  Process,  Large  New  Power  Plant Facility,  1400  MW
                        Process Flow  Diagram

-------
                           tf
                      13 Segments Each
                      9'-0' High
                      Catalytic Converter
Acid Tower & Mist
  Eliminator
                                               T-S' Dia. x OT-ff1
                                                Acid Cooler
 >. '** • '••I V
L' -NCFS
 MAL«
                                                                 	I	
                                                                    —f~—
                                                           J««J	
                                                              ^..t*wArij,.«,   r*T» •
                                                                                 MON»ANT<»
                                MRC-NAPCA PROCESS
                              LARGE NEW POWER PLANT FACILITY
                                      I40OMW
                                                                                                        PLOT PLAN
                    •ACILITY   V. ..'

                         ri—
                    ^^^^^1 T~Mkt
Figure 31*.   MRC-NAPCA Process,  Large  New Power Plant  Facility,  1*100 MW
                Plot Plan

-------
I
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                r

c
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                                                                                AcidTwwriMist
                                                                                  Eliminator
EcononHnr    Air Heater
                                                               30--01




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                                                    I.D. Fan
                                                         25'-0' —
         K-I
                                                  fj. ,• riH(0«!Sl SM'
                                                  D»I1S. .« ««< IN l«.C
                                                  TCIt»»NCtS
                                                  aCOMAt*     hfc
                                                   «» - »
                                                   «XX ,      AVOtfS
                                                   - V • - «•*if   . , .1C
                                                      t « fs • ,
                                                        C<.RI>OH/> nON
                                               DAV11IN LAIM>t«ATI*H%
                                  MRC-NAPCA PROCESS
                                LARGE NEW POWER PL ANT FACILITY
                                       1400 MW
ELEVATION
            Figure  35.   MRC-NAPCA Process,  Large New Power  Plant Facility, 1400  MW
                           Elevation Drawing

-------
                                                                     Table  8
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                                                              SO; REMOVAL FROM FLUE GAS



                                                       CATEGORY: Large New Power Plant Facility



                                                                  PROCESS FLOW SHEET
                         NAME OF PROCESS:  MRC-NAPCA Process
                                                                       MK:
                                                                              FLUE GAS  RATE:   2.5 x  106  SCFM
Process Stream
from boiler to converter.

Ib/hr
ll.SlxlO6
Flow
GPM SCFM
2.50X106
Temp.
°P
850-900
Composition
N2
70.27
CO;
21.66
H20
1.37
02
3.00
in Weight Percent
SO 2
0.61
NOx SO 3
0.05 	
HjSOfc CK<,
	 -— »
from converter to  economizer    11.8lxlOe


from economizer to air  heater   ll.SlxlO6


from air heater to acid tower   ll.SlxlO6


from acid tower to stack       11.67xl06
                                                                                                     1.22  2.86  Trace  0.05  Trace   0.;^


                                                                                                     4.22  2.86  Trace  0.05  Trace   0.?J
                    2.19xl06     865    70.27   21.66


                    2.t9xl06     665    70.27   21.66


                    2.')9xl06   150-500  70.27   21.66   1.22  2.86   Trace  0.05  Trace


                    2.17X106   200-225  71.07   21.91   3-98  2.89   Trace  0.05  Trace  Trace  	
from acid tower to storage
13.26x10"   195.67
                                                               100
25.00
                                                                                                                     75-00  	
o
Z

-------
                                                   Table  9
                                             SO;  REMOVAL FROM FLUE CAS
o
CA
z
o
TO
5
PI

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o
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TO
3
CATEGORY: Large New Power Plant Facility
PROCESS PLOW SHEET
NAME OP PROCESS: MRC-NAPCA Process
Process Stream
from boiler to converter
from converter to economizer
from economizer to air heater
from air heater to "tee"
from "tee" to acid tower
from acid tower to stack
from acid tower to storage

Ib/hr
2.95x10*
2.95x10*
2.95x10*
2.95x10*
1.18x10*
1.146x10*
1. 66x10*
Plow
0PM SCPM
62.5x10"
	 62.3x10"
	 62.3x10"
	 62.3x10" •
	 31. 2x10"
	 30.9x10"
20.146
Temp.
op
850-900
865
665
150-500
1450-500
200-225
100
FLUE QAS RATE:
62.5 x
10" SCFM
Composition In Weight Percent
70.27
70.27
70.27
70.27
70.27
71.07
	
C02
21.66
21.66
21.66
21.66
21.66
21.91
	
H20
4.37
1.22
1.22
1.22
1.22
3-98
25.00
02
3.00
2.86
2.86
2.86
2.86
2.89
—
S02
0.61
Trace
Trace
Trace
Trace
Trace
	
NOx
0.05
0.05
0.05
0.05
0.05
0.05
—
SO 3
-__
Trace
Trace
Trace
Trace
Trace
	
H2SO* CH%
	
0.81 	
0.81 	
0.81 	
0.81 	
Trace 	
75.00 	
NOTE:   Large stream of 2.5 x  10*  SCPM Is divided Into four  parallel streams


-------
 economizer and air  heater.   The  economizer  takes heat  from  the  flue
 gas  to  heat boiler  feed  water  while  an  air  preheater is used to
 recover more heat from the  flue  gas  by  heating the air prior to its
 entry into the boiler firebox.   Further cooling and condensation of
 sulfuric acid mist  is accomplished in packed acid towers which
 operate in conjunction with acid coolers.   Heat recovered in the
 acid coolers is used to  preheat  boiler  feed water.  The acid towers
 in this process, Figure  33,  are  reminiscent of the absorbers in a
 contact sulfuric acid plant.   In the latter, however,  the primary
 function is truly absorption,  i.e.,  of  S03  into slightly diluted
 sulfuric acid where more  sulfuric acid  is formed.  In  the case of
 the  flue gas process, sulfuric acid.exists  as such in  the gas and
 will condense as a mist  if  the gas is cooled.  Consequently, the
 main function of the acid towers here is to cool the flue gas to
 the  dew point of the acid which  then condenses into the liquid in
 the  tower.   The relationship of  SOa  in  the  flue gas to strength of
 the  acid condensing is ehowjTin  Appendix, 10 .to this volume.

 The  fine sulfuric acid mist not  retained in the acid tower  is removed
 by high efficiency  (about 95%) mist  eliminators prior  to discharge
 of the  flue gas to the stack.  Acid  produced by this process is
 stored  for  shipment as a 75 -  80$, or fertilizer grade, sulfuric
 acid.

 One  of  the  major decisions  in designing a process for  hanging onto
 the  end  of  a power plant is the  capacity for which to  design.  In
 this case,  we  have designed for  100% capacity all of the time,
 knowing, however, that this may  not be realistic.  On.the other
 hand, designing for a lesser capacity would mean sacrificing emission
 control  during  periods of peak capacity.  In Figure 36 a scheme is
 presented which permits the process to be run at incremental
 capacities  or at full capacity with only three of the  four  converters
 on line.  Each  converter has a 10? overdesign.  All four heat-
 recovery trains are available, at all times  to minimize loss of
 thermal  efficiency.

 Figure  35 is  an elevation view of one conversion train while Figure
 34 is a  plan  view.   Both figures give some  idea of the size of the
 treatment plant.

 Instrumentation for the MRC/NAPCA process is comparatively  simple
 (Figure  37).  Pressure drop through the catalytic converter  indicates
 fly ash  accumulation in the catalyst beds.  Instruments and control
 equipment maintain proper temperature,  pressure, draft and  liquid
 flow rates  to maximize treatment and power plant efficiency and
minimize interference with the power plant  operation.

Tables  10 through 14 present specifications for basic equipment
 employed in  this Case I process  design.
                                75


                   • MONSANTO RESEARCH CORPORATION •

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CT\
                                                                                        -*- VALVES

                                                                                        1 Catalytic Converters
                                                                                        2 Economizers
                                                                                        3 Air Heaters
                                                                                        4 Acid Towers and Mist E liminators
                                                                                        5 I.D. Fans
                                                                                        6 Stacks
                                               DECIMALS
                                                 Xk  - I
                                                 xt.*  i
                                                 »>I
-------
I
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2
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                                            B
     Flow Recorder
     Flow Recorder Controller
     Level Indicator
     Pressure Indicator
     Pressure Recorder
PDRC Press. Differential Recorder Controller
Tl   Temperature Indicator
TR   Temperature Recorder
     Temperature Recorder Controller
                   FR
                   FRC
                   LI
                   PI
                   PR
                   TRC
— Process Piping Lead
	Instrument Piping Lead
	I nstru merit E lectrical Lead
->v- Instrument Air Lead
 O Local Instrument
 B Board-mounted Instrument
 <8> Transmitter
 £ Motor Control Valve
 £ Hand Control Valve
                                                                                                                    TO
                                                                                                                   STACK
                           ELE
                                        UNUSS OTHF.BWISf SPIOFICO
                                        DIMENSIONS ARE I* INCHIS
                                        TOLERANCES
                                        DECIMALS      FRACTIONS
                                          »X = ±      t
                                          K XX - -      ANGLES
                                          XXXX BASIC      i 3O
                                        ALL SURFACES  Vf
                                                            MATERIAL

                                                            FINISH 	
                                                                                 CMtCKtD
                                                  MON3ANTH ResE.AH4.il f'(

                                                            l>AVJ*fS I Mti'KAltm

                                                               1-A '. Itt.-H L IIIIO
                                               MRC-NAPCA PROCESS
                                           LARGE NEW POWER PLANT fKtUTY
                                                     I40OMW
                                                                                                 SCALE «0(l«
                                                                                                                                INSTRUMENT FLOWSHEET
                                                                                                                               SHCl-T
                                            B
            Figure  37-
           MRC-NAPCA  Process,  Large  New  Power  Plant  Facility,   1400  MW
           Instrument  Flowsheet

-------
                             Table 10
                        CATALYTIC CONVERTER

1.  GENERAL
    Number of reactors in parallel - 4
    Number of catalyst beds in parallel in each reactor - 13
    Overall height of each reactor - 117 ft.
    Diameter of each reactor - 30 ft.
    Height of each segment - 9 ft.
    Height of each bed - 6-1/1 in.
    Size and shape of catalyst - 3/8 in., cylinders
    Total volume of catalyst - 20,000 cubic feet

2.  OPERATING CONDITIONS OF EACH REACTOR
    Total gas entering - 2.95 x 106 Ib/hr
    Inlet temperature - 850°P
    Outlet temperature - 865°P
    Operating pressure - Slightly below atmospheric
    Superficial gas velocity through each bed - 3-22 ft/sec.
    Pressure drop - 0.96 in. H20

3.  MATERIALS OF CONSTRUCTION
    Converter shell - Carbon steel
    Supports, beams, grates - Carbon steel (SA-285 Grade C)
    Blank plates between each segment - Cast iron
                                78

                   • MONSANTO RESEARCH CORPORATION •

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                             Table 11
              ACID RECOVERY AND MIST COLLECTION TOWER

1.  GENERAL
    Number of towers in parallel - 8
    Overall height of each tower - 70 ft.
    Diameter of each tower - 32 ft.
    Packed height In each tower - 16 ft.
    Size and type of packing - 3" Intalox Saddles

2.  OPERATING CONDITIONS OF EACH TOWER
    Total gas entering - 1.48 x 106 Ib/hr (31.2 x 10" SCFM)
    Total liquid entering - 1.94 * 106 Ib/hr (2,320 GPM)
    Gas inlet and outlet temperatures - 500°P and 225°F
    Liquid inlet and outlet temperatures - 100°P and 225°F
    Operating pressure - Slightly below atmospheric
    Pressure drop through the acid recovery tower - 8 in.
    Pressure drop through the mist eliminator - 10 in.
3.   MATERIALS OF CONSTRUCTION
    Tower - Carbon-steel vessel with minimum of 4" of acid-proof
            brick lining.
    Packing - Chemical stoneware
    Mist element - Chemically resistant glass fibers
                                79
                   • MONSANTO RESEARCH CORPORATION •

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                             Table 12
                            ACID COOLER
1.  GENERAL
    Size - 7' - 6" 0 x 20'
    Type - Shell and Tube
    Sq. Ft. Surface/Shell - 27,750
    Number of Units - 8
2.  PERFORMANCE OF ONE UNIT
    Fluid Circulated
    Total Fluid Entering

    Temperature In
    Temperature Out
    Pressure Drop
                                Shell Side
Water
1.125 x 106 Ib/hr
(2,250 GPM)
80°F
181°F
20 PSI
                         Tube Side
75% Sulfurlc Acid
1.9^ x 106 Ib/hr
(2,320 QPM)
100°F
225°F
10 PSI
          Heat Exchanged - 114 x 106 Btu/Hr
          MTD (Corrected) - 25.2°F
          Transfer Rate Service - 163 Btu/(Hr) (ft2)(°F)
3.   MATERIALS OF CONSTRUCTION
    Tubes:   Welded Carpenter 20
    Shell:   Carbon Steel
                                80
                   • MONSANTO RESEARCH CORPORATION •

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                             Table 13
                             ACID PUMP

1.  GENERAL
    Pump Type - Centrifugal
    Number of Units - 8

2.  OPERATING CONDITIONS OF EACH UNIT
    Liquid Pumped: 75% Sulfuric Acid
    Capacity, % 225°F, Normal 2,300 GPM, Design 2,540 0PM
    Specific Gravity % 225°F - 1.6?
    Viscosity § 225°P - 6 cps
    Diff. Head - 65 ft.

3.  MATERIALS OF CONSTRUCTION
    Casing - Stainless Steel, Alloy 20
    Impeller - Stainless Steel, Alloy 20
    Shaft - Stainless Steel, Alloy 20
    Packing - None

4.  DRIVER DATA
    Motor HP - 100, RPM - 1160
    Phase - 3, Cycle - 60,  Volts - 440
                                81

                   • MONSANTO RESEARCH CORPORATION •

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                             Table 14
                         INDUCED DRAFT PAN
1.   GENERAL
    Type of Fan - Double Inlet Centrifugal Fan
    Number of Units - 8
    Size: Height - 13 ft, Width - 18 ft, Length - 10 ft

2.  OPERATING CONDITIONS OF EACH UNIT
    Inlet Temperature - 225°F
    Suction Pressure - 30 In. Water
    Qas Density - 0.06 lb/ft3
    Capacity - 400,000 CFM % 225°F and 14.7 psla

3.  TYPE OF FLOW CONTROL
    Discharge Damper

4.  MATERIAL OF CONSTRUCTION
    Carbon Steel

5.  DRIVER DATA
    Motor HP - 3,000, RPM - 1,200
    Phase - 3, Cycle - 60,  Volts - 2,300
    Accessory - Brushless type exciter
                                82

                   • MONSANTO RESEARCH CORPORATION •

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In addition to the general assumptions previously cited, the
following assumptions were also employed in the Case I capital
and operating cost estimates:
        Flue Gas Rate

        Size of Power Plant

        Coal Required for
        Power Plant

        Operating Factor

        S02 Conversion

            Recovered as
        Flue Gas Temperature

        Cost of 3" Intalox
        Saddles
       2.5 MM SCFM

       1400 Mw

       580 tons/hr


       330 days/year § 100$ capacity

       90%

       95%


       850 to 900°F

       $4.55/ft3
Tables 15 through 18 show capital and operating cost estimates: for
Case I.  Figure 38 shows the effect of product sale price on net
operating cost.

C.  Case I-A.  Reversible Dry Absorption of SOg
This case differs from Case I primarily in the mode of product
recovery.  In this process, sulfur dioxide in flue gas is first
oxidized to sulfur trioxide over 'vanadia catalyst, as in Case I.
However, the product sulfur trioxide is removed from the main gas
stream by sorption on a dry medium, Na^Oi^, at relatively high
temperature.  Desorption is effected at still higher temperature
by decomposition of the sorption product.  The reactions involved
are:
S03
                                     850°F
                                            Na2S207
                               1000°F
                       Na2S207   =    Na2SOi« + S03
                                83
                   • MONSANTO RESEARCH CORPORATION •

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  30
£
3
**>
i
i
8.
4*
810
       Q.6
                                        Break Even Point
 Mills per Kilowatt Hour
 Ol2      !        0.2
                1.0
03        0       Oi5
   $ per Ton of Coal
1.0
15
     Figure  38.  Effect  of Product  Credit on Operating Cost  of
                  MRC/NAPCA Process  (Large New Power Plant Facility)
                     • MONSANTO RESEARCH CORPORATION •

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                        Table  15




               CAPITAL COST ESTIMATE SUMMARY



   Category:  NEW POWER PLANT (HIGH TEMPERATURE EFFLUENT)



                         Case HI





Name of Process: MRC-NAPCA Process  Flue Gas Rate;  2.5   MMSCFM

1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
21.
MW 1,100

Purchased Equipment
Purchased Equipment Installation
Piping
Instruments
Insulation
Electrical
Building
Land & Yard Improvements
Utilities
Incremental Cost of Precipitator
Acid Tower Packing
Mist Elements
Initial Catalyst Cost
Physical-Plant Cost
Engineering and Construction
Direct Plant Cost
Contingency & Contractor's Fee
Fixed Capital Cost
Capitalized Cost of Catalyst
Working Capital
TOTAL INVESTMENT
Capital Requirements
$/KW Capacity 21-71

Cost - $
6,632,000
2,539,000
2,653,000
995,000
800,000
663,000
663,000
995,000
995,000
1,500,000
155,000
2,3^0,000
630,000
21,860,000
1,372,000
26,232,000
5,216,100
31,178,100
2,110,000
1,021,000
31,639,100


% of Total
19.15
7.33
7.66
2.87
2.30
1.91
1.91
2.87
2.87
1.31
1.31
6.76
1.82
63.10
12.62
75.72
15-15
90.87
6.18
2.95
100.00

                               85
                 • MONSANTO RESEARCH CORPORATION •

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                                                 Table  16


                                       EQUIPMENT COST ESTIMATE  SUMMARY


                             Category:  NEW  POWER PLANT  (HIGH TEMPERATURE  EFFLUENT)

•                                                                       '  •

2                         Name of Process: HRC-NAPCA Process Flue Gas Rate: 2.5 MMSCPM
o

                                                  MU 1*00
m
CD
m

m                       Item _    No. of Units       Cost -
                        -    - —
5
o
I

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O
1.
2.
3.
H.
5.
6.
Catalytic Converters 4
Acid Towers & Mist
Acid Coolers
Acid Pumps
I.D. Fans
Storage Tanks
Eliminators




8
8
8
8
2
792,000
2,080,000
2,780,000
64,000
720,000
196,000
                                  Purchased Equipment Cost                        6,632,000

-------
                                                         Table 17


                                             WORKING CAPITAL  ESTIMATE SUMMARY
o
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Category: NEW POWER PLANT (HIGH
TEMPERATURE EFFLUENT)
Name of Process :MRC-NAPCA Process Flue Gas Rate: 2.5 MMSCFM

Item
1.
2.
3.
4.
5-
6.
7.
MW: 1400

Raw material inventory, 1 Month
Direct labor, 3 Months
Indirect cost, 3 Months
Operating supplies, 3 Months
Fixed costs
Spare parts
Miscellaneous

Cost - $
13,200
217,000
235,600
55,600
222,300
185,300
92,000

Percent
1.30
21.25
23.08
5.44
21.78
18.14
9.01
                                                         TOTAL
                                                                  1,021,000
100.00

-------
                             Table 18
                   OPERATING COST ESTIMATE SUMMARY
                 Basis:  330 Day/Year 6 100JE Capacity
         Category:   NEW POWER PLANT (HIGH TEMPERATURE EFFLUENT)
Name of Process MRC-NAPCA ProcessFlue Gas Rate
2.5
MMSCFM

1.
2.
3-
1.
5.
6.
7.
8.
9.
10.
11.
12.
13-
11.
15.
16.
17.
18.
19.
20.
21.
22.
23.
21.
MW 1,100
Fixed Capital Cost

ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15% of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20* of 2 & 3
Plant Overhead, 50? of 2, 3,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed
Capital/Yr.
Taxes, 255 of Fixed Capital
Insurance, 1% of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 1-78
Mill/kwh 0.7371
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mlll/kwh

31,178,100

TOTAL $
219,000
105,000
23,^00
1,573,900
236,090
895,700
_
3,083,090
25,700
969,200
_
_
_
991,900
3,117,810
62Q.S70
311,780
_
i.oqp.iqo
8,170,180




PER CENT
3-05
1.28
0.28
19-27
2.89
10.97
-
37.71
0.31
11.86
-
-
-
12.17
38.53
7.71
3-85
-
50.09
100.00



                      • MONSANTO RESEARCH CORPORATION •

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The desorbed sulfur trioxide is fed to an absorption tower
essentially the same as that in a contact sulfuric acid plant,
to make 100$ sulfuric acid and oleum.

As shown in the flow diagram in Figure 39, flue gas at high tempera-
ture  (850° to 900°F) flows through a catalytic converter where 90$
of the sulfur dioxide is converted to sulfur trioxide.  Next it
passes through an absorber-stripper where S03 is absorbed on a dry
medium consisting of porous silica or alumina granules impregnated
with  20% sodium sulfate.  Sulfur trioxide is stripped from exhausted
sorbent with hot gas from a furnace at 1000°F.  The absorber-
stripper is a fixed bed contactor designed very much like the
converter to minimize pressure drop through the absorbent beds.
Assuming the absorbent requirement is the stoichiometric amount,
each  absorber-stripper contains 13 bed sections in parallel and is
30 ft. in diameter by 117 ft. in height.

Four  absorption towers are required to permit adequate cycle, timing.
This  is illustrated by the following.table:


                          ABSORBER NUMBER
15 min abs.           	               	                	
15 min des.        15 min abs.          	                	
15 min des.        15 min des.       15 min abs.           	

15 min cool        15 min des.       15 min des.        15 min abs
Clean, hot gases emerging during the absorption cycle, at.about
820°F, are passed through an economizer and an air heater where
heat is recovered for utilization in the power plant.  An induced
draft fan moves the clean gas through the system and up the stack.

The hot gas emerging during the desorption cycle is cooled to about
500°F in a gas cooler.  Lower temperature would result In premature
condensation of acid.  The cooled gas is then directed to an acid
recovery unit for producing concentrated acid.  The sulfur trioxide
concentration in this gas stream is high, as shown by the material
balance in Table 19-  Thus, the primary function of the acid tower^,
in this case, is absorption of S03 into 96-98$ sulfuric acid.   Mist
eliminators are, again, required on the tail gas from the absorber.
The interesting point is that the gas rate to the absorber is  two
orders of magnitude less than the main gas stream flow.  Consequently,
only a single acid tower is required compared to eight in Case I.
                                89

                   • MONSANTO RESEARCH CORPORATION •

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                                                       Absorber-Stripper
                                               820>F
                                                     To   Hot    To
                                                    Boiler Water Boiler Air
620°F
                                                                                                      300°F
          Catalytic
          Converter
From
Electrostatic
Precipitator
LQ
                                                                                 • A.
                                                                                 r  1
                                                             '.»'*», -*r: •wn^.ii.
i.* ¥tNSf«'Ni «•'? .'« ',''•'.*)•?
rc;.m»N« i'-"
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MOVHA.Vr«.> Hi>-T. >i»t 1! t ••«!•«. HAT ION
-.H1..X , *, i: , •... . \
us \ t i • V •'
Reversible Dry Absorbent Process ! prorjK^ Fk» Dianram
Laroe New Power Plant Facility Process now oiagrai
1400 MW >
•••-•-• >-Nonej-v ;;, ( : i ^^j [V "CT"
         K-E -.-:.-.:
          Figure 39.   Reversible  Dry Absorbent  Process,  Process  Plow Diagram

-------
                                                                                 Table  19
                                                                           SO, REMOVAL FROM FLUE GAS


                                                                     CATEGORY: Large Hew Power Plant Facility


                                                                              PROCESS FLOW SHEET
                               NAME  OF  PROCESS: Reversible Dry Absorbent
                                                                                   MW: HOO
                                                                                      FLUE GAS RATE:   2.5 x 106 SCFK
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Process Stream
from electrostatic precipltator to
catalytic converter


from catalytic converter to absorber-
stripper


from absorber-stripper to economizer


from furnace to absorber-stripper


from absorber-stripper to gas cooler


from gas cooler to acid tower and
mist eliminator


from acid tower and mist eliminator

to I.D. fan


from economizer to air heater


from air heater to I.D. fan


from I.D. fan to stack


from acid tower to storage

Ib/hr
ll.SlxlO6
ll.SlxlO6
11.73xl06
O.llSxlO6
0.199xl06
0.199xl06
O.llSxlO6
11.73xl06
11.73xl06
11.817xl06
n i *i n&
Flow
GPM
2
2
	 2
2
3
3
2
2
2
	 2
i n? ii

SCFM
.50xl06
•50xl06
.igxio6
.5x10"
. 1x10"
.1x10"
.5x10"
.igxiu6
.I9xl06
•5xl06
Temp.
°F
850-900
865
820
1000
1000
150-500
225
620
300
300
i nr,
Composition
N,
70.27
70.18
70.69
70.18
11.60
11.60
70.19
70.69
70.69
70.69
C02
21.66
21.61
21.79
21.61
12.82
12.82
21.61
21.79
21.79
21.79
H2
1.
14.
1.
1.
2.
2.
1.
1.
1.
1.
n
0
37
10
15
10
61
61
10
15
15
15
t;
°2
3.00
2.93
2.96
3.07
1.82
1.82
3.08
2.96
2.96
2.95
in Weight Percent
S02
0.61
0.06
0.06
0.65
0.38
0.38
0.61
0.06
0.06
0.07
0.05
0.06
0.06
0.06
0.03
0.03
0.05
0.06
0.06
0.05
S03 KjSO,,
	 	
0.73 —
Trace 	
	
10.71 	
10.71 	
Trace 	
Trace 	
Trace .. 	
Trace 	
	 QQ ^

-------
This difference  is  reflected  in  the  comparative economics of  cases
I and  I-A.   Case I- A  still  requires  four  converters, but only one
acid tower.

Additional assumptions  employed  in the  cost estimates  for Case  I-A
are:

        Absorption  Efficiency  a)  100$,  i.e., stoichiometric,
                                   for  15 minutes
                               b)  50%  stoichiometric  for 15
                                   minutes

        Desorption  Efficiency  100$  in  30 minutes
      •  Amount of NazSOi* in     20$  (by weight)
        total absorbent

        Cost of Absorbent       5 shows the effect of product credit on net
operating cost.
                                92


                   • MONSANTO RESEARCH CORPORATION •

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        Table 20
CAPITAL COST ESTIMATE SUMMARY
Name
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
21.
Category: NEW POWER PLANT (HIGH
Case #1A
of Process: Reversible Dry Absorbent
MW moo
Purchased Equipment
Purchased Equipment Installation
Piping
Instruments
Insulation
Electrical
Building
Land & Yard Improvements
Utilities
Incremental Cost of Precipitator
Acid Tower Packing
Mist Elements
Initial Catalyst Cost
Physical-Plant Cost
Engineering and Construction
Direct Plant Cost
Contingency & Contractor's Pee
Fixed Capital Cost
Capitalized Cost of Catalyst
Working Capital
TOTAL INVESTMENT
Capital Requirements
$/KW Capacity 12-^8
TEMPERATURE EFFLUENT)
Flue Gas Rate: 2.5

Cost - $
2,934,000
1,400,000
1,470,000
440,000
353,000
294,000
294,000
440,000
440,000
1,500,000
30,000
	
"668,000
10,263,000
2,053,000
12,316,000
2,463,000
14,779,000
2,140,000
551,000
17,470,000

MMSCFM
% of Total
16.79
8.01
8.41
2.52
2.02
1.68
1.68
2.52
2.52
8.59
0.17
—
3.82
58.75
11.75
70.50
14.10 .
84.60
12.25
3,15
100.00

                  93
  • MONSANTO RESEARCH CORPORATION •

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                                              Table  21
                                    EQUIPMENT COST ESTIMATE SUMMARY
                          Category:  NEW POWER PLANT (HIGH TEMPERATURE EFFLUENT)
                    Name of Process: Reversible Dry Absorbent  Flue Gas Rate:  2.5  MMSCFM
                                                 MW 1400
•
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o                      Item	               Cost - $
a
w                      1.  Catalytic Converters                            792,000
£  vo                   2.  Absorber-Strippers                              792,000
o  •*=•
1                      3-  Furnace                                         200,000
§                      4.  Gas Cooler                                       65,000
o                      5.  Absorber and Mist Eliminator                    205,000
H                      6.  Acid Cooler                                      50,000
z                      7.  Regenerator Blower                               70,000
*                      8.  I.D. Fans                                       610,000
                       9.  Storage Tanks                                   150,000
                             PURCHASED EQUIPMENT. COST                   2,93^,000

-------
                                              Table  22


                                   WORKING CAPITAL ESTIMATE SUMMARY
o
z
z
H
O

3)
m
M
m
>   vo
3D   U1
O
Z

o
o
a
TJ
o
O
Z

ame

Category: NEW POWER PLANT (HIGH TEMPERATURE
of Process: Reversible Dry Absorbent Flue
MW: moo
Item
1.
2.
3-
4.
5.
6.
7.

Raw Materials & Chemicals Inventory f 1 Month
Direct Labor, 3 Months
Indirect Cost, 3 Months
Operating Supplies, 3 Months
Fixed Costs
Spare Parts
Miscellaneous
TOTAL
EFFLUENT)
Gas Rate: 2.5

Cost, $
70,000
113,000
115,600
24,300
97,100
81,000
50,000
551,000

MMSCFM

Percent
12.71
20.51
20,98
4.41
17.62
14.70
9.07
100.00

-------
                           Table  23
                 OPERATING COST ESTIMATE SUMMARY
                Basis:   330 Day/Year g 100$  Capacity
         Category:  NEW POWER PLANT (HIGH TEMPERATURE  EFFLUENT)
Name of Process: Reversible Dry Absorbent  Flue Gas Rate:  2.5 MMSCFM
                           MW 1100
                Fixed Capital Cost: $11.779.000
1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13-
15.
16.
17.
18.
19.
20.
21.
22.
23-
21.
ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 15? of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20$ of 2 & 3
Plant Overhead, 50$ of 2, 3,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10$ Fixed
CapHaT/Yr .
Taxes, 2$ of Fixed Capital
Insurance, 1$ of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 1.013
Mill/kwh U.132
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh
TOTAL $
731,000
me; nnn
23,100
739,000
110,800
615,000
	
2,351,200
25,700
189,100
	
_ —
	
511,800
M77,900
295, 6$0
117, 8QO
	
1,921,300
1,790,3o'o



PER CENT
15.26
2.1Q
0.19
15.13
2.31
13.16
___
19.15
0.51
10.21
— __
— — _
___
10.75
30.85
6.17
3-08
	
10.11
100.00



                                      96
                      • MONSANTO RESEARCH CORPORATION •

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   40-i
   30-
to
*h_
3
   20-
el-
's
                                              Break Even Point
8 10-
   0
              0.6
 0.4
Mills per Kilowatt Hour
   0.2
0.2
                                     0.6
              1.5
1.0
0.5
 $ per Ton of Coal
0.5
                               1.0
            Figure  40.  Effect  of Product Credit  on Operating
                         Cost  of Reversible Dry Absorbent Process
                         (Large  New Power  Plant Facility)
                                      97
                       • MONSANTO RESEARCH CORPORATION •

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«
§
I
5
i
3
   co
                                                                                    TO STACK
                                          FROM
                                      ELECTROSTATIC
                                      PRECIPITATOR
                                                     ACID TOWER
                                                      AND MIST
                                                      ELIMINATOR
                                                       474°F
                                                                     100°F
ACID
  ER
I.D.  FAN
   • SULFUR 1C ACID TO STORAGE
	WATER IN
     WATER OUT
                                                                       ACID
                                                                       PUMP
                                                                             COMBUSTION
                                                                               GAS




•jMi :* OTHfo«fl«f Vt^.r.ri;
p'vt -.^' -. , »wf is (Nt. <«r%
TO_H»M».-f »
ur"" V*LV. r-(»f:<-v*
• V
» 1 X ' B*«. ' . • JO
Al.l. »U •:'»*£* .
MATtB.A. 	 ....

«»»C |
»«
AVf-.. 1
»»o ~!
..^•.^T
- t ...&!•. f- -
	

	
t~ •— -
::•"»
MOXS-XNTO Kji.^rAIU tl
1 , 'X 1 M«t it »
. •. ' X. • I . >
MRC-NAPC A PROCESS
4MAI 1 FYRTMfi PtlU/FP PI AUT MTI ITV
220 MW
«.:.-,Fliae:v, -'.-
	 1 	 ~ 	 U__
( '• :!• I*«IIA I'|I>.N
PROCESS FUMf OM8RMM
»•. •>-. j«i»
S'ifF.I ;tf
           Figure 4l.   MRC-NAPCA Process, Small  Existing  Power Plant  Facility,  220 MW
                          Process Flow Diagram

-------
                                                                         Table
                                                                     30? REMOVAL FROM FLUE HAT-
2
o
z
.z
H
O
K
m
o
x

0
o
a
•o
o
    vo
CATEGORY:
NAME OP PROCESS: MRC-NAPCA Process
Process Stream
from boiler to flue gas heat exchanger
from flue gas heat exchanger to burner
from burner to catalytic converter
from catalytic converter to flue gas heat exchanger
from flue gas heat exchanger to acid tower
from acid tower to stack
combustion gas to burner
product acid to storage


2.
2.
2.
2.
2.
2.
0.
26

Ib/hr
36x10'
36x10'
58x10'
58x10'
58x10'
56x10'
22x10'
.5xl03
Snail Existing Power Plant Facility
PROCESS FLOW SHEET
HW: 220 FLUE GAS RATE: 0.5
Plow
GPM
	 0.
0.
	 0.
	 0.
	 0.
	 0.
	 0.
39.13

SCFM
50x10'
50x10'
55x10'
55x10'
55x10'
55x10'
05x10'
Temp.
op
300
683
850
865
171
225
100
x 10' SCFM
Composition in Weight Percent
N2
70
70
70
70
70
71
73
CO 2
.27
.27
.52
.52
.52
.25
.20
21
21
20
20
20
21
_
.66
.66
.87
.87
.87
.09
	
H2O
1.37
1.37
1.69
1.71
1.71
1.53
25.00
O2 S02
3
3
3
2
2
2
_
.00
.00
.08
.95
.95
-98
22

0.
0.
0.
0.
0.
61
61
59
06
06
Trace
_
	
N'0X SO 3
0.05 	
0.05 	
0.05 	
0.05 0.03
0.05 0.03
O.on Trace
—
h2SOi, CKw
	 	
___
	 	
0.77
0.77
I
Trace 	
75.00 	
O
z

-------
                        Table  25



               CAPITAL COST ESTIMATE SUMMARY



           Category:  SMALL EXISTING POWER PLANT



                         Case #2





Name of Process:  MRC-NAPCA Process  Flue Gas  Rate:   0.5  MMSCFM




                        MW  220

1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
21.
Purchased Equipment
Purchased Equipment Installation
Piping
Instruments
Insulation
Electrical
Building
Land & Yard Improvements
Utilities
Incremental Cost of Precipltator
Acid Tower Packing
Mist Elements
Initial Catalyst Cost
Physical-Plant Cost
Engineering and Construction
Direct Plant Cost
Contingency & Contractor's Fee
Fixed Capital Cost
Capitalized Cost of Catalyst
Working Capital
TOTAL INVESTMENT
Capital Requirements
$/KW Capacity 16.12
Cost - $
2,011,000
770,500
1,007,000
302,100
211,700
201,100
201,100
302,100
302,100
500,000
51,000
520,000
126,100
6,512,700
1,308,600
7,851,300
1,570,300
9,121,600
126,000
365,200
in ?i? ftnn

% of Total
19.73
7.55
9.87
2.95
2.37
1.98
.1.98
2.95
2.95
1.89
0.52
5.09
1.23
61.06
12.81
76.87
15.38
92.25
1.18
3.57
100 00

                               100
                •  MONSANTO RESEARCH CORPORATION

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o
z
H
m
CO
                                                   Table  26

                                        EQUIPMENT COST ESTIMATE SUMMARY
                                     Category:  SMALL EXISTING POWER PLANT
                          Name of Process:  MRC-NAPCA Process    Flue Gas Rate:  0.5 MMSCFM

                                                     MW  220
                                                       -

                      Item                                             Cost -  $
a  2                  Catalytic Converter                              278,000
o  *~^
1                     Acid Tower & Mist Eliminator                     340,000
°                     Acid Cooler                                      540,000
o                     Acid Pump                                         12,500
H                     Flue Gas Heat Exchanger                          595,000
z                     I. D. Fan                                        180,000
*                     Flue Gas Burner                                    8,500
                      Storage Tank                                      60,000
                                     Purchased Equipment Cost        2,014,000

-------
                                                     Table 27



                                         WORKING CAPITAL ESTIMATE SUMMARY



                                       Category: SMALL EXISTING POWER PLANT



                       Name of Process:   MRC-NAPCA Process   Flue Gas Rate:    0.5  MMSCFM



*                                                 MW 220

z
w

z
-I
o
a                       Item	Cost -  $	        Percent
m


>  o
21  ro
O
X

o
o
3)


a

H

O
•z
1.
2.
3.
"•
5.
6.
7.
Raw Material Inventory
1 Month
Direct Labor, 3 Months
Indirect Cost, 3 Months
Operating Supplies
3 Months
Fixed Costs
Spare Parts
Miscellaneous
36,200
78,600
79,700
16,500
66,000
55,000
33,200
9.92
21.52
21.82
4.51
18.07
15.06
9.10
                                             Total     365,200               100.00

-------
                            Table  28
                   OPERATING COST ESTIMATE  SUMMARY
                  Basis:   330 Day/Year € 60jT Capacity
                 Category:   SMALL EXISTING  POWER PLANT
Name of Process MRC-NAPCA  Process Flue  Gas  Rate
0.5
           MMSCFM

1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
15-
16.
17.
18.
19.
20.
21.
22.
23-
24.
MW 220
Fixed Capital Cost 9

ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5% of Fixed Capital
Supplies, 151 of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20$ of 2 & 3
Plant Overhead, 50? of 2, 3,
4 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, IQ % Fixed
Capital/Yr.
Taxes, 2% of Fixed Capital
Insurance, 1$ of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
COST: $/Ton of Coal 6.51
Mill/kwh 2.66
BY-PRODUCT CREDIT
ADJUSTED OPERATING COST
ADJUSTED COST: $/Ton of
Coal
Mill/kwh

,421,600

TOTAL $
444,250
79,000
15,600
471,080
70,700
141,000
.-
1,221,630
18,920
318,190
-
-
-
337,110
942,160
188,440
94,220
_
1,224,820
?)7fl^,'i60




PER CENT
15.96
2.84
0.56
16.93
2.53
5.07
-
43.89
0.67
11.44
-
-
-
12.11
33.85
6.77
3-38
-
44.00
100.00



                                     103
                     • MONSANTO RESEARCH CORPORATION •

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  60i
  50-
  40-

   30
Q_
"o
C
o
8
Profit
     0.5
                  Break Even Point
Cost
                   Mills per Kilowatt Hour
                0,5       1,0       1,5
                                                 3.0
      To
              1.0     2.0     3.0
                    $ per Ton of Coal
                          4.0     5.0     6.0     7.0
    Figure 42.   Effect  of Product Credit  on Operating Cost of
                  MRC^NAPCA Process (Small  Existing Power  Plant
                  Facility)
                     • MONSANTO RESEARCH CORPORATION •

-------
Additional assumptions employed in material balance  and  cost.
estimates for Case  II are:

      .  Flue Gas Rate          0.5 MM SCPM

      •  Size of Power Plant    220 Mw

      •  Coal Required for      90 tons/hr
        Power Plant

      •  Operating Factor       330 days/year @ 60% capacity

      •  SO 2 Conversion         90%

      •  SO 3 Recovered as       95%
      •  Flue Gas Temperature   300°F

      •  Cost of 3" Intalox     $I».55/ft3
        Saddles

      •  Cost of Natural  Gas    35
-------
in height, and one acid tower, 20 ft. in diameter by 46 ft. in
height.  The converter consists of nine catalyst beds in parallel.

For this case, the pertinent assumptions, over and above the general
ones cited earlier, are as follows:
        Off-Gas Rate

        Copper Production
        Capacity

        Operating Factor

        SOa Conversion

        303 Recovered as
        Off-Gas Temperature

        Cost of 3" Intalox
        Saddles
90,800 SCFM

230 tons/day


330 days/year § 100? capacity

90$

95$
$4.55/ft3
Here, as in all previous cases, the amount of vanadia catalyst is
derived from Figure 32.  Tables 30 through 33 show capital and
operating cost estimates and Figure 4*1 shows the effect of product
sale price on net operating cost.
                                106
                   • MONSANTO RESEARCH CORPORATION •

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S   3   ,
                                         B

                                                                      Acid Tower
                                                                       and Mist
                                                                      Eliminator
            Offgas
Electrostatic    Heat
Precipitator   Exchanger
Offgases
  From
Smelting
 Plant
                                                                                 Add
                                                                                Cooler
                                                        450°F
                            850°F
                                                                                                             REVISIONS
                                                                                         '•W
                                                       Sulfuric Acid to Storage
                                                       Water In
                                                       Water Out
                                                                                          Catalytic
                                                                                          Convwter
                                                        TOLCB»KCCS

                                                          KX -- J*       '
                                                          xxx -- *       ANUI t:
                                                          xr »* BA?'C   -  »»o'
                                                        •ii  si^r»cE» i/
                                                        m»rc» »L
                                                                                                   MtlNSANfl-
                                                                                                            r i > . -• :. »IM tti v :\.»i v
                                                   MRC-NAPCA PROCESS
                                                      SMELTER BJCILTTY
                                                                                         PROCESS FU
           Figure  43.    MRC-NAPCA  Process, Smelter  Facility

-------
                                                                             Table  29
Z


o
m
(A


5
x
O
X

o
o
a
•o
o
5
     CD
                                                                        S02 REMOVAL FROM FLUE CAS





                                                                       CATEGORY:  Smelter Facility




                                                                           PROCESS PLOW SHEET
NAME OF PROCESS: MRC-NAPCA Process
Flow
Process Stream
from smelting plant to electrostatic
•* precipltator
from electrostatic precipltator to
heat exchanger
from heat exchanger to converter
from converter to heat exchanger
from heat exchanger to acid tower
from acid tower to stack
product acid to storage
Ib/hr
0.13x10*
0.13x10*
0.13x10*
0.13x10*
0.13x10*
0.11x10*
21,600
0PM SCFM
	 90
90
	 90
	 89
	 89
	 85
36.3 -
,800
,800
,800
,081
,081
,913
—
Temp.
op
150
150
850
918
523
218
100
Composition
N2
75.25
75.25
75.25
75.25
75.25
79.77
	
C02
5-36
5.36
5-36
5-36
5.36
5.68
	
H20
5.61
5.61
5.61
1.85
«.85
3.61
25.00
02
10.50
10.50
10.50
9.77
9-77
10.36
	
FLUE
GAS RATE: 90,800
SCFM
in Weight Percent
SO 2
3.25
3.25
3.25
Trace
Trace
Trace
	
NOx SO 3 M2SO»
	 	 	
	 	 	
	 Trace 1.25
	 Trace 1.25
— Trace Trace
	 	 75.00
CH»
	
	
	
	
	
	

-------
                         Table 30




               CAPITAL COST ESTIMATE SUMMARY



                Category:  SMELTER FACILITY.



                         Case # 3




Name of Process: MRC-NAPCA Process   Off-gas Rate; 90.800 SCFM

1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
•13.
11.
15.
16.
17.
18.
19.
20.
Purchased Equipment
Purchased Equipment Installation
Piping
Instruments
Insulation
Electrical
Building
Land & Yard Improvements
Utilities
Incremental Cost of Precipitator
Acid Tower Packing
Mist Elements
Initial Catalyst Cost
Physical-Plant Cost
Engineering and Construction
Direct Plant Cost
Contingency & Contractor's Pee
Fixed Capital Cost
Capitalized Cost of Catalyst
Working Capital
TOTAL INVESTMENT
Cost - $
1,306,000
524,700
653,000'
195,900
156,800
130,6.00
130,600
195,900
195,900
'
12,000
226,000
164,000
3, 891, 400
778,300
4,669,700
934,000
5,603,700
553,500
214,730
6,371,930
% of Total
20.50
8.24
. 10.25
3.08
2.46
2.04
2.04
3.08
3.08
__
0.18
3-55
2.57
61.07
12.21
73.28
14.66
87.94
8.69
3.37
100.00
                                 109
                   •  MONSANTO RESEARCH CORPORATION •

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o
z
M
                                                       Table 31


                                            EQUIPMENT  COST  ESTIMATE SUMMARY


                                              Category:   SMELTER FACILITY


                             Name  of Process:  MRC-NAPCA  Process  Flue Gas  Rate:  90.800 SCFM
a
m
at

>  £                    Item                                      Cost  -$
n  o
I

o                       Electrostatic  Precipitator                 ^30,000
o
a)
3                       Catalytic  Converter                       1^3,000
a

H                       Acid Tower & Mist  Eliminator               320,000
o

1                       Acid Cooler                               '96,000


                        Acid Pump                                    2,000


                        I.D. Fan                                   ^5,000


                        Off-gas Heat Exchanger                    205,000


                        Storage Tank                               65,000
                            Purchased Equipment  Cost             1,306,000

-------
o
z


z
H
O

70
m
u>
m
>
a
o
i

o
o
O
OJ
O
Z
                                                  Table 32


                                      WORKING CAPITAL  ESTIMATE  SUMMARY


                                         Category:  SMELTER  FACILITY


                       Name of  Process: MRC-NAPCA Process Off Gas  Rate:  90,800 SCFM
Item
1.
2.
3.
4.
5.
6.
7.
Raw Material Inventory, 1 Month
Direct Labor, 3 Months
Indirect Cost, 3 Months
Operating Supplies, 3 Months
Fixed Costs
Spare Parts
Miscellaneous
Cost - $
3
52
52
10
12
35
19
,420
,050
,200
,500 .
,030
,030
,500
Percent
1
24
. 24
4
19
16
9
.60
.24
.30
.88
.58
.32
.08
                             TOTAL                        214.730               100.00

-------
         Table  33
OPERATING COST ESTIMATE SUMMARY
Basis: 330 Day/Year g 100$ .Capacity
Category: SMELTER FACILITY
Name of Process MRC-NAPCA Process Off-ga§ Rate 90,800 SCFM

1.
2.
3.
1.
5.
6.
7.
8.
9.
10.
11.
12.
13.
11.
15.
16.
17.
18.
19.
20.
Fixed Capital Cost

ITEM
Raw Materials & Chemicals
Direct Labor
Supervision
Maintenance, 5$ of Fixed Capital
Supplies, 15$ of Maintenance
Utilities
Other
TOTAL DIRECT COST
Payroll Burden, 20$ of 2 & 3
Plant Overhead, 50$ of 2, 3,
1 and 5
Pack & Ship
Waste Disposal
Other
TOTAL INDIRECT COST
Depreciation, 10 % Fixed
Capital/Yr.
Taxes, 2$ of Fixed Capital
Insurance, 1$ of Fixed Capital
Other
TOTAL FIXED COST
TOTAL OPERATING COST
5,603,700

TOTAL $
in ,onn
R?,^no
1R.600
280,200
1?,000
18,^00
__
1lQr600
-,-, finn
195.200
__
__
— —
208.800
560,100
112,100
56,000
—
728,500
1.386.900

PER CENT
2.96
3.79
1.12
20.21
3.03
1.31
— _
_^2.A2_
0.98
11.07
—
—
—
15.05
10.11
8.09
1.03
—
52.53
100.00
                  112
     • MONSANTO RESEARCH CORPORATION •

-------
0.
 15
                                  Break Even Point
  U)       0.5       0        03
      Net Operating Cost, 10*Dollars/Year
1.0
1.5
Figure
                     Effect of  Product Credit  on
                     Operating  Cost of MRC-NAPCA
                     Process  (Smelter Facility)
                           113
             • MONSANTO RESEARCH CORPORATION

-------
                             REFERENCES
1.   Arthur G. McKee and Company, Systems Study for Control of
    Emissions - Primary Nonferrous Smelting Industry, Vol. Ill
    (Final Report), Contract PH 86-65-85, Plate No. C-l,
    San Francisco, June 1969.
                   • MONSANTO RESEARCH CORPORATION •

-------
          APPENDIX I
 COMMERCIAL CATALYST TEST DATA
              115





• MONSANTO RESEARCH CORPORATION •

-------
100
                          CATALYST -A-
                      10                 2:0

                    W/F, gin-sec/mole SO2><10~6
    Figure  45.   W/F-Conversion Profile for Catalyst
                 'A' at  900°F
                            116
              ,• MONSANTO RESEARCH CORPORATION •

-------
100
 10
                            CATALYST -A-
                     1.0                 2.0
                    W/F, gm.»ec/mole SO2><10~6
3.0
    Figure  46.  W/F-Conversion Profile for Catalyst
                 'A' at  850°F
                            117

              • MONSANTO RESEARCH CORPORATION •

-------
  100i
   90-
   80-
o
iu
   70
z
o
u
o
m
Z
iu
U
   5(
                                  Catalyst A
   40-
   30-
   20-
   10-
                         1.0                 2.0
                        W/F, gm-sec/mole SO2xlO~6
                                                                3.0
       Figure ^1.   W/P-Conversion Profile for  Catalyst
                     'A'  at 800°F
                                118
                 • MONSANTO RESEARCH CORPORATION •

-------
          Table 3^
   CATALYST  -'ATTEST RESULTS
Run
No.
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
* S02
in feed
(mole )
0.34
0.34
0.31
0.30
0.34
0.319
0.346
0.280
0.281
0.296
0.300
0.292
0.318
0.282
0.276
0.308
0.302
0.291
0.284
0.322
Catalyst
Weight
(g)
6.0
6.0.
6.0
6.0
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
Total
Flow
(cc/sec )
140
no
60
60
30
30
30
40
50
60
50
40
70
30
50
40
30
60
70
80
W/F
(g-sec/cc)
0.15
0.15
0.10
0.00
0.283
0.283
0.283
0.212
0.170
0.142
0.170
0.212
0,121
0.283
0.170
0.212
0.283
0.142
0.121
0.106
W/F
(g-sec/
mole S02)
1.07
1.07
0.787
0.81
2.03
2.16
2.00
1.85
1.48
1.17
1.38
1.77
0.93
2.45
I1. 50
1.68
2.29
1.19
1.04
0.80
Temp.
(°F)
800
850
850
900
850
850
850
850
850
800
800
800
800
800
900
900
900
900
900
900
%
Conversion
52
68
68
68
83
88
89
98
90
49
59
75
45
90
87
88.
91
85
80
75
              119
• MONSANTO RESEARCH CORPORATION •

-------
10
                           CATALYST  B
                            " T~  T
                                                   _jj_
                1.0                2.0

                W/F, gm-**c/mol« SOj


Figure  48.   Effect of W/F  on Conversion with
             Catalyst-B at  900°F
                                                          3.0
                           120
              •  MONSANTO RESEARCH CORPORATION •

-------
100i
                            CATALYST B
                     1.0
 1
2.0
3.0
                    W/F, gm-iec/mole SOj xlO~
    Figure  49..  Effect of  W/F on Conversion  with
                 Catalyst-B at 850°F
                            121
              • MONSANTO RESEARCH CORPORATION •

-------
10
                          CATALYST  B
                     1.0
                        I
                       2.0
 i
3.0
                    W/F, gm-»«c/mol« SO, x10~
   Figure 50.
Effect of  W/F on Conversion  with
Catalyst-B at 800°F
                            122
              • MONSANTO RESEARCH CORPORATION •

-------
           Table 35



     CATALYST-B  TEST  RESULTS
Run
No.
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
% SOp
In feed
(mole)
.300
.292
.302
.286
.325
.302
.293
.285
.301
.302
.322
.305
.309
.295
.302
.307
.304
.314
Catalyst
Weight
(gms)
8.5
8.5.
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
8.5
Total
Flow
(cc/sec)
30
40
50
60
70
80
30
40
50
60
70
80
30
40
50
60
70
80
W/P
^ cc '
.283
.213
.170
.142
.121
.106
.283
.213
.170
.142
.121
.106
.283
.213
.170
.142
.121
.106
W/P
/ gin-sec ^
^mole SCV
2.30
1.78
1.38
1.21
.91
.86
2.36
1.82
1.37
1.15
.92
.85
2.24
1.76
1.37
1.13
.97
..82
Temp.
800
800
800
800
800
800
850
850
850
850
850
850
900
900
900
900
900
900
Conver-
sion
80.2
75.3
73.4
69-1
66.5
60.6
89-0
83.5
80.9
73.9
71.5
69.0
89.0
87.0
85.2
82.7
80.0
77.1
              123
• MONSANTO RESEARCH CORPORATION •

-------
lOO-i
 90-
                              CATALYST  E
 10
                      1.0
 i
2.0
 i
3.0
                     W/F, gm-s»c/moU SO2
     Figure 51.  Effect  of W/F on Conversion with
                  Catalyst-E at 900°F
                             124
               • MONSANTO RESEARCH CORPORATION •

-------
                    CATALYST E
                                                    3.0
               W/F, gm-sec/mole SO2 X10"
Figure 52.   Effect  of W/F on Conversion  with
             Catalyst-E at 850°F
                       125
         • MONSANTO RESEARCH CORPORATION •

-------
lOOi
                            CATALYST E
                                                          3.0
                     W/F, gm-»ec/moU SOj
       Figure 53.  Effect  of  W/F on Conversion with
                   Catalyst-E at 800°F
                              126
                • MONSANTO RESEARCH CORPORATION •

-------
          Table 36
    CATALYST-E TEST RESULTS
Run
No.
1
2
3
1
5
6
7
8
9
10
11
12
13
11
15
16
17
18
In feid
(mole)
.31*
.299
.301
.303
.309
.297
.311
.318
2.0
.315
.280
.315
.299
.296
.312
.296
.309
.311
Catalyst
Weight
(gms)
8.5
8.5
8.5
8.5
8.5
8.5
5.0
8.5
. 8.5
8.5
.5.0
5.0
8.5
8.5
8.5
8.5
8.5
8.5
Total
Flow
(cc/sec)
30
50
10
60
30
10
10
50.
60
60
60
80 .
30
10
50
60
70
80
W/F
cc
.283
.170
.213
.112
.283
.213
.125
.170
.112 .
.112
.083
.063
.283
.213
.170
.112
.121
.106
W/F
, gm-sec »
vmole S00
2.20
1.39
1.73
l.ll
2.23
1.75
.97
1.19
.17
1.10
.726
.18
2.31
1.76
1.33
1.17
.96
.82
Temp.
800
800
800
800
850
850
850
850
850
850
850
850
900
900
900
900
900
900
Conver-
sion
65.3
19.2
59.5
10.9
92.7
85.7
72.6
78.6
16.0
77.1
56.0
11.0
91.0
88.7
86.3
81.9
77.0
72.9
              127
• MONSANTO RESEARCH CORPORATION •

-------
                                                     Table  37
Catalyst
  code
             Catalyst   Typs of Support
                                         Support   •   trtl
                               Promoter  HattrlH  Catalyit
                                                                       Ota
                                                                       Plow
                                                               tttf     Rat*      S02
                                                             Pronoter  oo/tto  Cone. <
                                  Convorcion
                                  Efficiency
                                      I
                                                           75*F
Mli.l-l
NKb-J
MEC.-l
MEOb-l

MKBC-1
rfECV-2
MtAV-1
MECV-2
KtHV-1
ME'-;V-2
MCrV-2
MENV-1
MlJPV-2
MKKV-6
HENV-?
MENV-;
t5»DIi-10
«E3-2
«EPV-»
MEPV-4
MENV-1
MEPV-2
HECV-2
MESV-1.
ME7V-S
KBLV-1
MENV-1
NSKN-1
MEKV-]
MELV-1
MEKV-1
MEKM-2
MSV-1
KEPV.t
WSPF-1
rtKPV-1
MEBV-1
MEKV-1
SnOj
HaO
Cl-,0,
P»0
SnOj
Cr:0,
BaO
VjO,
VjOj
VjO.
           Fuller"o Earth
                         Fuller's Earth
VjOj.pcjO,  Fuller's Earth
V;05           "
             VjO,
             V20S
v,o,-re2o,
VjOj
V;0j
v,o,
                         Puller'* Earth
             S«02
V,0b
Vj05
v,05
Vj05-Pe,0,
V,0,-PejO,
V,05-Pe,0s
            AljOj (Acid)
            AljOj (AJll)
            Fuller1! Earth
V,05
MnOj
V,05
             V,05
 V,0S
 MnOj
 V,0,
              v,o,
              FejO,
              VjOj
            Puller's Earth
            FiilJer'n Eurth
            Fuller's Earth
            Fuller's Earth
            Fuller's barth
            Fuller'« Earth
                         Puller's Earth
           Puller's  Earth
           Alumina
           Puller's  Earth
                 (acid)
                        Puller'o Lurth
Non«






Sec,


CuSO,
SO.
KSCN
NaSCN
SeOa

KjCO,
K,COj
HaSCN
NaSCN
CuSO,
SeO,

NaSC.'l
L1HSO,
NaSCH
SeO,

KHSO
KHSO

L1HSC

KHSO,
KHSO,
SeO,
K,CO,

P«l(SO,),
KHSO,

BljO,
KHSO,
92
92
92

92
92
92
72
150'F

72
72
72
72
72
200'F
72
72
225'F
72
250'F
72
72
300'F
72
72
72
72
72
o?
9*
350*F
72
72
72
72
72
72
100'F
72
72
72
72
f»50'F
72
72
175'F
72
500'F
72
72
72
72
550*F
72
72

88.
72
8
0
8
4
4
H
4
a

\
. a
a
•a
a


a
8

e

a
8

8
a
8
8
8

a
a
a
8
a
8

a
8
8
8

a
8

8

a
a
a
8

a
D
a
8
a
                                                                             20
                                                                20
                                                                20
                                                                20
                                                                20
                                                                             20
                                                                             20
                                                                             20
                                                                             ?0
20
20
20
                                                                20
                                                                20
                                                                             20
                                                                        to
                                                                        10
                                                                        10
                                                                        40
                                                                        to
                                                                        to
                                                                        40
                                                                        to
        to
        to
        40
        to
                                                                        40
                                                                        to
                                                                                     to
                                                                                     1*0
to
to
to
to
                                                                         40
                                                                         to

                                                                         to
                                                                         to
               0.317
               O.)12
               0.292
               0.295

               0.293

               0.281
      0.300
      0.301
      0.301
      0.300
      0.303
                                                                               0.312
                                                                               0.306
                                                                                            0.320
20
20
20
20
20
20
20

20
20
20
20
20
20
20
20
20
20
20
20
to
to
to
to
to
to
to
40
to
tu
tc
to
to .
to
40
to
to-
to
to
to
0.311
0.301
0.310
0.312
0.324
0.307
0.311
0.319
0.298
0.298
0.308
0.307
0.298
0.298
0.3C3
0.315
0.302
0.297
0.301
0.309
                                                                                            0.316
                                                                                 0.313
                                                                                 0.282
                                                                                 0.798
                                                                                 0.298
                  .
                 0.278

                 0.?8*
                 C.310
                 I.t9
                 J.t<
                 1.57
                 1.18

                 l.tc

                 0.23
                                                                                          l.Cl
                                                                                          1.60
                                                                                          1.60
                                                                                          1.61
                                                                                          0.22
                          1.56
                          1.68
                                                                                                       1.6t
                                                                                          1.57
                                                                                          1.69
                                                                                          1.57
                                                                                          1.55
                                                                                          1.62
                                                                                          1.67
                                                                                          0.22
                                                                                          1.35
                           .
                          1.66
                          1.59
                          1.62
                          l.ta
                          1.61
                          1.77
                          1.70
                          0.23
                          1.66
                          1.88
                                                                                                       1.77
1.85
1.95
I.t8
I.t9
                  1.91

                  1.63
                 .1.63
                                                                                                    19.2*
                                                                                                    19.?•
                                                                                                    lf.l«
                                                                                                    14.6*

                                                                                                    lt.0«

                                                                                                     2.1
         22.3
         lt.0
          B.6
          2.0
          0.3
                            1.3
                            O.b
                                                                                                                  1.3
                                                                                                     1.6
                                                                                                     0.9
                                     2.6
                                     2.5
                                     2.2

                                     5:1
                                     0.0
                           t-i.6
                           18.1
                            7.3
                            5.5
                            1.0
                            0.0
                            7.3
                            1.9
                            1.7
                            1.1
                            1.3
                            1.0
                                                                                                                  1.3
1.3
1.1
 .0
o.o
                                                                                                    i.t
                                                                                                    i.t
                                                                                                    1.3
                                                       130
                                  •  MONSANTO RESEARCH CORPORATION •

-------
                                          Table  37  (Cont'd)
Catalyst
  CoJa
                    Tyca of Support
                      Fuller1g Earth
                                   Oae
          »'»                      Plow
         Support     wtl       Htf    Rate      SO,
Promoter  Material  Catalyst  Promotar  oe/aac  Concf
                                                                   	      40      0.323
                                                                                                Convorelon
                                                                                                   3.1
                                                  600'F
rltC>'-2X3 Cr-Ba-Sn 	 	 .— - 	 .
MiiPV-'j

.VitHV-S
,*ii-PF-l
Mafv-3

H ^-69- *' ' 1
HliKV-2
MEPV-3
MEBC-2

Mt.RV-lq
MEBC-3

HEKV-12
MERV-l-K
600
514CP-4
"•IERV-1P
MhRV- 1 5'

?t
V205

CuO
Bao
V205
CuO
BaO
v2os
V20,

Pt-Fe
•1,0.,
V205
Pt-Fe
Pt
V20,
CuO
V20,
Pt-Fe
V205

v2os5
V205
V205
V205
ZnO
MoOj
V205
V205
V;0.
V205
V205
V205
V;c5
Fe203
V205
CuO
Sn02
Fe203
V205
v,o5
V205
MoO 3
MoOj
Fe20,
Cr20j
Fc203
CuO
V20,
V:05
V205
WOj
Pe203
ZnO
V205
ZnO
Fe»0 1
MoO,
V20,
Pe203
CuO
CuO
CrjOj
CuO
r>nO 2
v2os
Ta,0,
CuO
Cr203
Vj05
Pe,03
Kn02
V205
V20,

B2ol
Fe203
V26V
V205
Puller's

it
•i
n

Alumina
Puller's
"
n

"
n

n
"

Alunlna
Puller's
fl
Alumina
Alumina
Puller's
n

Alumina
Puller's
' "
"
11
"
"
11
"
"
"
"
n
"
"
"
"
n
n

n
it
"
-7

N
n

n

n
n
™

n

"

n

"
"

11

n

ii
n
n

•

n
"
"
"

n
SK-JJ10
Puller's
Earth .

"
"
n


Earth
n
n

•n
n

n
**


Earth
11


Earth
**


Earth
n
n
it
n
n
n
"
•t
"
"
**
n
it
n
n
n
"

n
n
H
n

0
n

"

ti
n
it

n

n

n

n
n

n

n

n
n
"

n

"
n
"
"

"

Earth
K2C03
KHSOW
KHSO,.
Potash
K2C03
KHSO,.
...V
KHSO,
KHSO,
None

Rb2SO,
KHSO,

C8-Rb-K
Rb2SO,
K2SO,
.. —
RbjSO,
Rb 2SO,

Pe2(SO,,)3
Rb-Cs-K
KHSO,

.._.
Cs-Rb
Potash
KHSO,
Ca2SOv
06,30,
Cs-Rb-K
Potash
Potash
KHSO,
Cs 2SO,
Cs-RbrK
Rb2SO,
C3-R6-K
C8-Rb-K
Pe2(SO,)3

Cs2SOi.
None


/Te2(SO|,)3
K2SO,


....
KHSO,

	

Cs2SO,
Na2$o,
	

	

	



Ll2SOi.
KHSO,

....

KHSO, '

LlHSOk
Na2beO,
KHSO,

	

KHSO,
Bl,0)
KHSO,
....

KHSO,
KHSO, '
K2C03
. 72

72
72
72

....
72
72
84

72
64

72
76

.. —
72
72
. 	

72
64

....
72
72
76
72
72
72
72
72
72
72
72
72
72
72
72

72
84

80
72
72
84

80
64

84

72
72
84

84

84

84

72
64

84

64

7;
72
64

84

72
88
72
84

64
72
72
8

8
8
8

.._.
8
6
8

8
8

8
8


8
8
	
	
8
8

__ — .
8
8
4
8
8
8
8
8
8
8
8
8
8
6
4
4
8
8

20
8
8
8
8
20
8

8

8
8
8
8
8
8-
D

8

8
8

8
8
8

a
8
'8

.8
8
8
8
8
8

8
8
8 '

10
10
10
20
10
10
	
20
20
None
20
20
20
8
8

20
20

....
20
20
• _..
20
20
20
20
20
20
20 .
20
20
20
20
20
20
20
20
20
Hone
-.__-
20
20
	
....
20
	
20
20
	
— -
	
	
20
20
	
?0
20
20
20
	
....
4
20
	
20
20
20 '
40
40

40
40
40

40
40
.40
40
40
40
40
40

' 40
10
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
• 40
40
40
40
40
40
40
40
40
40
40
40
40
40
40
0.301
0.316

0.316
0-303
0.309

0.301
0.298
0.309
0.307
;0.315
0.307
• 323
.288

0.303
0.31?
0.313
0.301
0.299
0.313
0.317
0.303
• 334
0.295
0.301
0.334
0.334
0-335
0.301
0.303
0.299
0.318
0.334
0.312
0.319
0.301
0.300
0.331
0.323
0.307
0.305
0.312
0.300
0.308
0.293
0.295
0.324
0.321
0.297
0.302
0.317
0.302
0.324
0.290
0.294
0.319
0.295
0.319
0.293
O.J93
0.296
. 0.292
0.295
0.315
0.292
0.309

1.62
1.87

1.87
1.52
1.31

1.64
1.74
1-31
1.59
1.84
1.71
1.41
1.58-

1.53' .
1 .86
1.85
1.22
. 1.29
1.45
1.50 .
1.05
1.36
1.59
1.56
1.53

l.:J6
1.62
1.53
1.69
1.46
1.36
1.32
1.42
1.51
1.67
1-75.
1.47
1.49
1.41
1.55
1.60
•1.51
1.72
1.54
1.79
1.38
1.58
1.60
1.59
1.63
1.29
1.71
1.69
1.73
1.89
1.46
1.80
1.64
1.69
1.59
1.96
1.34
1.78
1.72

92.0*
75.1

46.2
38.6*
34.3

34.2*
30.2
20.1*
20.0*
15.1*
15. 0»
10.8
9.0

8.9
8.6
8.5.
8.0
8.0
7.2
6.9
6.9
6.5
6.4
6.3
5-7
5-1
5.1
5.1
5.0
5.0
4.6
4.5
. 4.5
4.4
4.3
4.0
•3.9
3.7
3-6
3-6
3-6
3-3
3.2
3-1
3.1.
3-1
3.1
3.0
3-0
2.8
2.6
2.5
2.3
2.0
'1.9
1.7
1.6
1.4
1.4
1.4
1.0
1.0
0.0
0.0
0.0

                                                     131
                                 • MONSANTO  RESEARCH CORPORATION •

-------
     Table  37  (Cont'd)
Cui.alyst
Coilo Cotulyjt

'•liiCI'-J
KLFV- 1

IIP-1
MERV-l-K-600
MENV-1
511-CP-1
MEKV-12
MEKV- 7-600
MECV-3-K-600
MEKV- 12-600
jll-CP-3
MEKV-8-600
511-C?-2
MEKV-11-600
MKKV-9-600
MEKV-10-600
MEFP-1
MEPV-1

MEPV-3
MEST-I
MEKV-1
MEPV-1
MEKV-2
MEKV-1 '
KELV-1
CiEyv - 1

MEKV- 'J
MEMV-1

MEKV- 3
MEP-1
MEZV-1

NKM-1
ME KM-?
MS2P-J

MEMP-1

HEPV-5
MEWV-1

ME9V-1
MEKV-5
H£BV-:

MEV-1
(TKKM-1

KKCV-3
HlilV-U

Ol'lCI-1
H<-
r't t,,*p 1
MUNP-1

514CF-2
514CP-3
NKKV-5
MEPP-1
MLPV-4
ME! V-l
MKPV-3
MfcKV-3
MEKV-1
MiCKV-4
MfcKV-1
MELV- I
MEKV- 5
MKZP-1
MF.P- 1
MLMF- 1

MfcUV- 1
MEKH-2


I'l-Pc
Pt
V,0,

Pt
V20S
Pt
Pt
V,05
V20S
V20,
VjO,
Pt
V205
Pt
v,o,
vloj
Pe,0j
V205
Pe,0,

Ta'jO,

vjoj

v|oj
V,0,
VjOj
Pe,0,

vao.
MoO)
pj,3,
V205
ZnO
MoO,
Mn02
Pe20,
ZnO
Pe,0,
MoOj
V20,

WO 3*
V20,

vjot
B20,
V205
KnOj

VjO.

Pt
Pe
rt
Pt
Pr
rt
PI
pt
V;,05
v'jGs'
vCo°'
VJOj
V;0 J
v,ob
viol
V,05
v,o5!
V,05
FejOJ
ZnO
Ke 20 '
MoO]

WnT
V Oe
•/nO
Type of Support
Alumina
Alumina
Puller's

Alumina
Puller's
Alumina
Puller's
Puller's
Puller's
(I
It

''
«
"
It
"
"
11
n


n

II
"

"
Alumina
Puller's

n

n
n

n
n
"

"
w



Earth

Earth
Earth
n
n
Earth
Earth
n
n
n

"
«
w .
H
It
n
"
n


it

;
n

"

Earth

11

n
n

n
"
"

"


Puller's Earth
Al?0j 'acid )

Alumina
„
n
n
	 	
Fuller' a
11
n
n
it
„
ii
"
»
„
"

n
Alumina
Puller* 3







Earth
"
» '
"
„
"
I
«
„
«

n

Earth

Support xtl
Promoter Material Catalyst

».
Potash

Rb-K
NaSCN
Ce-Rb-K
Cg-Rb-K
ca-K
C8-Rb-K
CB-Hb-K
C0-Rb-K
Cs-Rb-K
C,-Rb-K
PejtSOJj

Pej(SOi)j
• Na»SeO»
KH30k
Potash


tiy30 1,



None

None
None

None

None

Nonf

None
None

B120,

Nor.e

None
KHSO^

CCjSO,
K2CO,

	
P«2(SO»))

NaSCN

KHSO,.
?ej(SO, >j
Potash
KhSOu
KliSO,
KHSO,
KIISO.
KHSO,,
L1HSO,
None
None
None
None

BIjCj

None

650'F
..-
72
700'F
z
72
72
72
72
72
72
72
72
72
7S

72
72
76
72
74
72
72
72


64

72
80
81

80
72
81

81

81
81

88
72
81

92
72
750'F
7?
72
800'F
	
	
__._
	
	 	
72
72
. 72
72
72
74
72
72
76
•11
72
81
81
80
84

88
72
84


	
a

8
8
8
8
8
8~
"e"
8
8
8
4
4
8
8
4
1
6
8
e
4
^

a
8
e
20
8
8
20
8
8

8
8
8
8

a
8
a

a
8

8
8

	
	
_-._
	
	 „_
8
8
4
a
8
6
8
e
4
4
4
8
8
8
8
8
20
3
8
8
8
8
8
Das
Clou
wtS Kate
Promoter oc/sec

—
20

20
20
20
20
20
20
20
20
20
20
20

20
20
20
20
20
20
20
20


....

20
---_

	
20
	

	

	
	

4
20
	

---.
20

2
2

	
	
	
	
	 •
20
20
20
20
20
iO
20
20
20
?P
20
	
	
	

4
20
?0

10.
no :
40

10
10
40
40
10
10
40
10
10
40
10
10
10
10
10
10

40
40
10
40
40
10
40
10


40

40
10
40

10
10
10

10

10
10

HO
10
40

40
40

40
40

10
40
"0
10
40
40
10
40
40
10
40
40
40
4U
40
40
40
40
40
'"0
40
40

40
40
40

SO]
cone. %
0.300
0.289
0.315

0.310
0.311
0.303
0.288
o'nt,
0.311
0.313
0.287
C.317
0.291
0.3H
0.327
0.331
0.312
0.321

0.310
0.288
0.306
0.300
0.320
0.312
0.307
0.311


0.310

0.302
0.307
0.305

0.301
0.299
0.312

0.309

0.311
0.303

0.302
0 . 309
0.316

0.305
0.301

0.297
0.298

0.290
0.2')5
0.292
0.306
0.305
0.296
0.308
0.309
0.324
0.291
0.313
0.202
0.301
0.289
O.J89
0.289
0.293
0.301
0.319
0.292
0.306
0.317

0.313
0.^16
0.302

Conversion
Efficiency
W/P t
2.30
1.77
1.49

1.11
1.33
1.91
1.61
1.11
1-31
1.32
1.45
1.28
1.31
1.08
1.33
1.39
1.36
1.60
1.56

1.39
1.61
1.42
1.56
1.62
1.65
1.62
1.70


1.55

1.55
1.19
1.66

i.53
1.69
1.56

1.60

1.55
1.55

1.51
1.72
1.35

1.42
1.66

1.57
1.49

1.7!)
1.01
1.32
1.91
1.52
1.07
1.19
1.72
1.51
1.72
i.50
1.48
\'.6?
L .00
I. SO
1.81
i .',:
1.66
l.so
l.Sfr

l.lt .
J.df)
1.17

10.7
37.0
3.8

33.5
16.5
15-5
12.8
11.2
10.7
10.6
10.3
8 7
8.5
8.2
6.9
6.6
6.5
6.1
5.6

5.5
5.2
l.l
3.5
•3-1
2.9
2.6
2.3

2.3
2.3

2.3
2.0
2.0

2.0
2.0
1.9

1.9

. 1.6
1.3

1.3
1.0
0.9

0.7
0.3

19.5'
15.1

80.3
6Q.4
63.0
60.1
27.9
27.0
26.3
10.7
E.6
7.9«
7.4
6.5
6.3
6.?
5.9
1.9
4.1
4.0
J-5
3.1
2.6
2.5

1.9
1.9
1.7

              132
• MONSANTO RESEARCH CORPORATION •

-------
                                           Table  37'  (Cont'd)
Catalyst
  Code
                       Type of Support
                                        vt*
                                      Support
                             Promoter  Material
                                                               Conversion
                                                               Efficiency
                                                                  it
MliWV-l

MEV-1
MEKM-1
MEM-1
MEBV-2

MEHV-1
V205
WO,
V205
Mn02
MoOj
V205
B203
V20S
Mod 3
None

None
XHSO,,
None
None

None
81

92
12
80
81

81
20
            20


            20
10

10
10
10
10

10
0.301

0.320
0.301
0.321
0.295

0.300
          1.56
.36
 65
,11
 15
                                                                                                1.60
1.6
1.6
1.5
1.1

1.3
                                                      850°F
MEV-l
             V205
                       Fuller's Earth
                                                     92
                                                                                10
                                                                                       0.315
                                                                                                 1.38
                                                                                                         10.5
                                                       900°?
illiPV-5
HP-1
MEPV-1-500
MENP-1
511CP-3
511CP-2
HI-:PV-S
MEPV-1
MERV-1-600
MtiRV-le
MECV-3a
MERV-lf
5HCP-1
MERV-lg
MliCV-3b
MECV-3
MERV-1-100
HLPV-3
HECV-3-100
MEPV-1
MECV-3-500
KEHV-1-500
HEPZ-1
MLRV-1
MECV-3C
WEPM-1
MEPV-3
MECV-3-600
MECV-3
HF.RV-1
MEKV-5
MEPV-ld
HERV-1-500
MEPZ-J
HEKV-j
Hf.KV-2
MliKV-1
MEPV-1
MUFF-1
KtiFV-1

HEf.V-1
Mlil.'V- 3
MEFV-3
ML.LV-1
HEFV- 1

KELV-2
MhK-1
liiPV-S

MkiZF-1

MEMP-1

MKKM-1
MliWV- 1

MKHV-1

MFV-1
HEUV-1
MEZV-1

NEKM-2
KEM-1
MEUV-;;

V205
Pt
v2o5
pt
rt
Pt
V20S
V205
V205
V20S
V205
V205
pt
v2o5
V20S
V205
V205
V205
V205
V205
V?05
v2os
ZnO
V205
V205
MoOj
v,os
V205
V205
V205
V205
V205
V205
ZnO
V205
Vj05
v2o5
v2o5
Fe203
V205
Fe203
V205
V205
v2os
v2os
V20S
Fe203
V205
Fc203
V20S
Fe203
Fe20j
ZnO
Fe20j
Mo03
MnO,
V20,
WO,
v2o5
Mo03
V205
v2o5
v2o.
ZnO
MnO,
MoOj
V20b
B203
Al203(acld)

Al203(acid)
Alumina
— --
.___
Al203(acld)
H
Fuller's Earth

n ti
n n
	
Puller's Earth
n 11
n ti .
n ii
n n
n tt
Al203(acld)
Puller's Earth

n ti
n ii
n n
0 II
n ti
ii n
n n
ii «
SK-IUO
Ai2o3(acid)
Fuller's Earth
ii it
n it
» n
n n
n n
. u n
" "

n ii
n n
n rt
n »
tt 11

n . n
n ii
rr ir

n ' u
..

u ii
tt n

tr n
1.
II II
II il
Alumina
Fuller's fc'arth


KHSO.*
pp. / Cf), \ .
r e2 \ oui^ ; 3
KjCOj
NaSCN
. 	
	
• KHSO,.
K2COj

Rb2SOi.
CsjSO,.

	
Rb2SOi.
Cs2SOi,
Cs2SOu
Rb2SO,,
KHSO,,
Cs2SOi,
K2C03
Cs250i,
Rb2SO,.
Potash
Rb2SOi»
Cs2SO,,
Potash
KHSO,,
Cs2SOi,
Cg-SOl.
Rb2SOi,
KHSO,,
K2C03
Rb2SOk
Potash
KHSO,.
KHSOi,
KHSO,,
Potash

F«!(So!j|

KHSOt,
Ha2SO,.
Fe2(SO,,)3
LiHSO,.
KHSOi.

Ll2SOi.
None
None

None
None

KHSO,.
None

None
None
B120,
None
KH.SOi,
None
None

72

72
• _— --
	
	
72
72
72
72
72
72

72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
72
71
76
72
72
72

72
72
72
72
72

72
80
81

81
81

72
81

81
92
88
8"
72
80
81

8

8
-—_-
	
	
a
8
8
8
8
a

8
8
8
8
8
8
8
8
8
8
8
8
8
8
8
8
8
8
B
8
8
6
6
1
8
8
1
1
B
8
8
8
• 1
1
8
20
a
8
8
Q
0
8
8
8
8
a
8
8
8
8
8
20
8
8
10

20
--_-
	
__--
10
20
20
20
20
20

20
20.
20
20
10
20
20
20
20
20
20
20
20
10
20
20
20
20
20
20
20
20
20
20
20
20
20

20
20
20
20
'20

20
	 	
	

	
	 	

20


	
	
1
	
20
„ 	
	

10
10
10
10
10
10
10
10
10
10
10
10
10
10'
10
10
10
10
10
10
.10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10
10

10
10
10
10
10

10
10
10

10
10

10
10

10
10
10
to
10
10
10

0.320
0 2QS
V * cy?
0.293
0.291
0.302
0.303
0.320
0.279
0.298
0.291 •
0.287
0.296
0.302
0.298
0.286
0.295
0.291 •
0.296
0.287
0.279
0.286
0.296
0.302
0.320
0.297
0.299
0.296
0.297
0.318
0.320
0.309
0.293
0.297
0.302
0.287
0.301
0.286
0.296
0.299
0.301

0.301
0.322
0.299
0.306
0.297

0.318
0.295
D.303

0.302
0,300

0.293
0.301

' 0.303
0.302
0.298
0.302
0.293
0.297
0.299

1.81
1.11
1.53
1.99
1.21
1.05
1.81
1.60
1.95
1.98
2.02
1.96
1.51
1.95
2.03
1.58
1.98
1.37
2.02
1.6o
2.03
1.96
1.62
1.29
1.96
1.55
1.37
1.96
1.16
1.29
1.72
1.55
1.39
1.62
1.63
1.57
1.52
1.58 .
1.67
1.66

1.90
1.38
1.11
1.83
1.78

1.31
1.56
1.53

1.60
1.61

1.71
l.'jl

1.59
1.11
1.56
1.68
1.72
1-57
•1.13

82.7
75 o,
II-6
68.0
67.1
67.7
65.6
62.1«
56.7
51. 8*
53.3
53.0"
19.7
' 16. 1»
11.1*
13.7
13.7
"3.2
13.0
11.8
38.9
37.1
36. l»
. • 35.6
33.3-
31. »
31.1
30.3
30.2
30.0
29.9
25.1
25.2
15.6
13.9
12.8
12.2
11. a
11.1
10. ?

9.9

g'o
7.2
6.1

6.3
5.8
5.6

5.0
1.7

3.6
3.0

3.0
3-0
3.0
2.7
2.7
1.7
1.3

                                                       133
                                   MONSANTO  RESEARCH  CORPORATION

-------
  EFFECT OF CATALYST PARTICLE GEOMETRY ON PRESSURE DROP IN A FIXED BED
In selecting the geometrical form for catalyst particles in a fixed-
bed reactor, several factors require balancing.  It is desired to
minimize pressure drop through the bed within the limits allowed by
contact time requirement and gas velocity, minimize catalyst attrition
rate, and minimize catalyst cost (i.e., not increase cost due to
unusual geometry).

A computer program was written to calculate pressure drop as an aid
in the design of the catalytic reactor.  Three equations, were included
in the program based on the magnitude of the Reynolds number.  For
laminar flow through the catalyst bed (Re <10.0) the equation becomes
(Ref.  1).
                AP -   200.0 GyLA2  (l.Q-0)2
                &* -        g Dp< p  6'


The transient flow (Re  10.0- 200.0) equation is  (Ref. 96):
                      2.0 f Gn LX3-n   (1.0-S)3"n              (2)

                              D(3-n> g  p 63
where f and n are evaluated from a least squares fit of two graphs
(Ref. 2).  of Re versus f and Re versus n.

The turbulent flow (Re > 200.0) equation is (Ref.  3.),.
                AP =  "-vg-o G1'9 u0'1 X1'1 (1.0-6)              (3)
                              g Dp1'1 p 63


Each of the above is then converted to psi/ft by:


                Ap =  P/11U.O
                                  136


                     • MONSANTO RESEARCH CORPORATION •

-------
                                 NOTATION

A    =    surface area of a pellet  (sq ft)
D_   =    effective particle diameter  (ft)
G    =    mass velocity of gas  (Ib/sq ft hr)
L    =    length of bed (ft)
Re   =    modified Reynolds number  (Dp up/y)
V    =    volume of a pellet (cu.ft)
f    =    modified friction factor
g    =    acceleration due to gravity  (4.18 x 108 ft/hr2)
n    =    state of flow factor
u    =    linear velocity of fluid  (ft/sec)
6    =    fractional voids in bed
AP   =    pressure drop (Ib/sq in/ft)
X    =    shape factor
y    =    fluid viscosity (Ib/hr/ft)
p    =    fluid density (Ib/cu ft)
For the catalytic reactor of this report, all Reynolds numbers were
in the turbulent flow range and equation (3) was used.
Three shapes were considered, those being, cylindrical, spherical,
and granular.  For the cylindrical case, the area and volume were
found for three lengths and diameters were L/D = 1.0 (D = 1/8,
2/8, 3/8 in.) and the following calculations were made:
                                          (Ref.  1)
               Dp =  (6VA)1/3
                                   137
                     •  MONSANTO RESEARCH CORPORATION •

-------
                         6 = 0.33


For the spherical case X = 1.0; Dp = 1/8, 2/8, 3/8 In.; and 6 <= 0.36.
For the granular case X = 1.0/0.7 (Ref. §8); 6 = 0.4; and DD = O.i5,
0.35, 0.525 in.

Figure  5^  summarizes the results of the .study of the effect of catalyst
particle geometry on pressure drop.  Based on these comparative results,
the suggested catalyst geometry is a cylinder about 3/8 inch in diameter
with a length-to-diameter (L/D) ratio of. 1 or 2.  This is a common
form for commercial catalysts..  A spherical shape would result in lower
pressure drop for equivalent, diameter, but it is more costly to. pro-
duce because of its shape.  The granular shape, with still lower Ap
response, is subject to higher attrition rate.
                                  138
                     • MONSANTO RESEARCH CORPORATION •

-------
   200
   160-
2
CD
U.
O

I
u
c

I
O
£
Ul
   126
             Cylinder  I/I
             C/llmUr  2/1
             Cylinder  I/I
              QruwUr  e.li

              OmwUr  o.J
   100 -
    25
                          10        16        20
                        FLUE GAS VELOCITY, U, Ft/toe
       Figure  54.   Effect  of Catalyst Particle  Geometry
                     on Pressure Drop in  a Fixed  Bed
                                139
                   • MONSANTO RESEARCH CORPORATION •

-------
                             Table  38
 »10CS(CARDtll3? PRINTER)
 • ONF WORD 1NTFGF.RS
 •LIST SOURCE PROGRAM
 C	
~C"	~"PROfiR"RM"TO~OnLCUT,ATE~ MKfSSURt DKUP  IN A CATALYTIC BED.
 C        PROGRAM IS IN FORTRAN-IV FOR THE  IBM 1130 BK  WITH DISK.
 C        THE EQUATION USED DEPENDS ON WHETHER THE GAS  FLOW IS
 C        LAMINEHt TRANSIENTt OR TURBULENT.
 C        THE REQUIRED PARAMETERS FOR THE EQUATIONS ARE-
 C        EFFECTIVE DlAv-ETER OF THE PELLET(DIAP)	
"C     	THE SHAPE"Mt IUK ur Iht HtLLti(ELMDA)
          THE VOID SPACE OF THE REACTOR USING  THESE PELLETS  (DELTA)
          THE GAS DENSITY (RHOI
          THE r,AS VISCOSITY (EMU I
          THE VELOCITY OF THE GAS '(U)
          DP IS PRESSURE DROP IN LB./ SO.  IN./ FT.AND  IS PER UNIT
          WFBSuxE "i>ii»T "f"tk'""K
       DIMENSION TITLEI3.6I
       SG=ft.l
       AL'UO
       DO 1 1=1.3                                      "  '
    1   READ12 .1010) (T 1TLE I I iJ) «J=1.6)
  1010 FORMAT (6A* I                                 .......
 C
------- -on -TO- j^rrj ---- ........ ------------------- : -
       WRITFIl.tlOXTITLEUtKI.K-l .6)
    10 FORMATllOX. 'PARAMETRIC STUDY OF PRESSURE DROP THROUGH CATALYTIC BE
      10
       (70 TO |20»21» 22>tJ
 C        DEFINE THE REOUIRFD PARAMETERS FOR THE CYLINDRICAL CASE
 C                •                   '                        '
    20 CONTINUF
       OELTA=0.33   •                   ......................
       HNa( 1.0/a. 01/12. 0
       TIAP=(6.0*V/3.142)«»0.3331
       El.MDA»0.205«A/V»»0.6667
       GO TO 23
          DEFINE THE REQUIRED PARAMETERS FOR THE SPHERICAL CASE
    21 CONTINUE
       ELMDA=1.0
       DELTA=0.36
       GO TO 23
 C                     .
 C        DEFINE THE REQUIRED PARAMETERS FOR THE GRANULAR CASE
 C                •               ......
    22 CONTINUE
-- OCLTA-0** -
       OIAP=0.15/12.0
                                      140
                   •  MONSANTO RESEARCH CORPORATION •

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                         Table  38  (Cont'd)
 c
    23 WRITEI3.1111>
  1111 FORMAT(lOX,'EFF OTA I FT I1.'  LTWDJT^     ',*  DtlTA
       WRITEI3.1110)I>IAP,ELMDA.DELTA
  1110 FORMAT!  10X»E12.5»1X,F7.4,6X»F7.4)
 C
 C        DEFINE GAS DENSITY AND VISCOSITY
 C
       RHO«U.U3Z5
       F.MU=0.0798
       WRITE(3«1211)RHO
  1211 FORMATdOXt'DENSHY OF THE GAS IS1.Ell.4,' LBS/CU FT')
       WR1TE(3,1212IEI»!U
       FORMATUOX.'VISCOSITY OF THE GAS IS',Eli.4,« LBS/HR/FT'/)
          DEFINE GAS VELOCITY
          CALCULATE THE MASS FLOWTGl
 C               i
       G=RHO»U»3600.0
       WRITE(3.1001)U
  1001 FORMAT1  lOX.'GAS VELOCITY (U)  ISSFS.Of' FT/S'E'C1)'
 C      	.	
~C        CTCfULAIt RtYNOCDS IMUMBEH [RE i
                      FVU »3600.
 C
 C        IF REYNOLDS NUMBER IS LESS THAN 10.0 USE THE FIRST EQUATION.'
 C _      IF RFYNOLOS NUMBER IS GREATER THAN  10.0 BUT LESS THEN 200.J
"C~"       CISE" THE bErDHn~tOUATION.
 C        IF REYNOLDS NUMBER IS GREATER THAN 200*0 USE THE THIRD EQUATION
 C                                                           .......
       IFIRE-10.0130*40,40
    30 CONTINUE
       WRITE I 1 • lUlllRt
   100 FORMATf  10X, 'LAMINAR FLOW INDICATED BY REYNOLDS NUMBER OF'tE12.5l
       OP»200.0»G»EMU«AL«ELMDA«»2»(1.0-OELTA)»*2 ...........
       DP»DP/ISG»OIAP»»2»RHO»DELTA«»3I
       OP=DP/144.0                                  ...........
       r;TT TTJ""90
 C
    40 IF(RE-200.0)50.50i60                                  "' .........
 C
    50 CONTINUE                                             ............
       WRITE! 3,2 00 1 RE _     _ _ _
   ZOO" I- UKlAll   rtJH", • I KAIN 5 1 TTDrn«L hLUW INDICAltD IJY KEYNOLD5 NUMBER QF'«E
      112. SI
       REMOD=ALOGTIRE>                        "            ........
       EN-0.522»REMOD*0.4B187
       F=10.0»»J-0.62«REMOD+1.6428)            .      ..........
       OP«2.0«F*G««EN«AL«ELMOA«»(3«0-ENI«(1.0-OELTAI**<3.0-EN»
       ,,^-(J(l/-|uJnt,ww | 3,u- tIM I w3
       OP = DP/144.0
       GO TO  90
                                      141
                    • MONSANTO  RESEARCH CORPORATION •

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                       Table 38  (Cont'd)
c
   60 CONTINUE
  300 FOHMATI   10X . MUKBULENT FLOW INDICATED BY REYNOLDS  NUMBER OF'.EIZ.
     15)
      (>P»0.02«.3»G»»1.9«EMU»»0.1»ELMDA»«1.1«I 1.0-DELTA)
      nP=nP/tDIAP»»i.l«SG»RHO«DELTA«»3)        	
   90 CONTINUE
     • WRITE(3idOO)DP
  40C FORMAT)   10X,   'OEL P«'tE13.5t' PSI/FT'/I
C
      U=U+10.0
   80 CONTT17OT
      WRITEO.11121
 1112 FORMAT I MM
   70 CONTINUE
      CALL  EXIT
      END
                                   112
                     'MONSANTO RESEARCH CORPORATION  •

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                      APPENDIX III REFERENCES
1.   Leva, M. , et al, "Introduction to Fluidization," Chemical
    Engineering Progress, Vol. M, No. 7, p. 518.

2.   Ibid, p. 519-

3.   Foust, A. S., et al, "Principles of Unit Operations," J. Wiley
    and Sons, Inc., New York, 1962, p. ^76.
                   • MONSANTO RESEARCH CORPORATION •

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                         APPENDIX IV
                       REACTOR  DESIGN
Preceding page blank
               • MONSANTO RESEARCH CORPORATION •

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                             APPENDIX IV

                           REACTOR DESIGN


Optimum design of the catalytic reactor was based on minimum total
capita- and operating.costs, which are functions of the costs of the
reactor, fan, and power calculated on a yearly basis.

Total Cost of Reactor  =  annual cost of (reactor+fan+power)

The total catalyst .volume was fixed by the. reaction rate, determined
experimentally, for SQ% conversion of sulfur dioxide to sulfur tri-
oxide.  Consequently, for a given reactor diameter, the height of
the catalyst bed is prescribed.  It then becomes a matter of determin-
ing the most economic geometry for the requisite catalyst volume.
If, for example, a given reactor diameter, or cross sectional area
of the catalyst bed is increased, the installed cost of the reactor
increases; but, due to a corresponding reduction in bed height, the
pressure drop decreases, reducing the cost of fan and power.  An
economic balance shows that at. one particular equivalent 'diameter,
the reactor cost is minimum.  An increase beyond this point raises
cost of the reactor much faster than it lowers the cost of fan and
power.  Hence, this represents the point of optimum reactor design.

The assumption employed in the optimization program for the large
new power plant facility were:

   1)  Flue gas rate:  2.5 MM SCFM

   2)  Flue gas temperature:  850°F

   3)  Total volume of catalyst:  20,000 ft3

   4)  Power cost:  $0.006/KWH

       Installed cost of reactor:  $2,500/ft (for 30-ft diameter)
5)

6)
       Installed cost of fan in $:  85.8 (x°<699), where x = drive
       horsepower
   7)  Depreciation rate for reactor and fan:  10 years, straight
       line depreciation

The design parameters for .the reactor installation for a 1400 MW sta-
tion, based on our optimization study are:

   1)  Number of reactors in parallel:  4

   2)  Number of .catalyst beds in parallel in each reactor:  13
                                 1116
                    • MONSANTO RESEARCH CORPORATION •

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   3)  Height of each bed:  6-1/1 in.
   4)  Height of each segment:  9 ft.
   5)  Overall height of each reactor:  117 ft.
   6)  Diameter of each reactor:  30 ft.
   7)  Size and shape of catalyst:  3/8 in. cylinders
   8)  Total pressure drop.:  0.96 in. H20
   9)  Total cross sectional area of catalyst:  38,000 sq.  ft.
  10)  Superficial gas-velocity through each bed:   3.22 ft/sec.
The computer program for optimum reactor design is  shown in Table  39
                                  1*17
                    • MONSANTO RESEARCH CORPORATION •

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                                Table  39
    PROGRAM OOPT - CALCULATES  THE OPTIMUM DIAMETER OF A  REACTOR  BED.
    REAL KW
    DIMENSION ZI20) .INI 10)
    PI-3.1M6

    INITIALIZATION  OF PARAMETERS FOR BED

 10 CALL READP(6tZ»IN>
    J=IN<1 I
    CICWZI 1)
    TOTV=Z(2)
    D=Z<3)
    DD=Z15)

    INITIALIZATION OF GAS PARAMETERS

 20 CALL READPIbtZtlNI
    TF=Z(1)
    EMU=Z<2)
    GM=Z(3)
    AMWZKtl
    P-ZI5)
    TR'TF+460.
    RHO-AMW«P/(10.T2«TRI
    1-0

    PRINT  INITIAL PARAMETERS

 30 WRITE! 3.600) J.CHW.TOTV.D. DO. THETA.RHO>EMUtGM(TF»AMWiP

    WR1TEI3.510)
 35
    H"TOTV/XS£Ct
    GUX»GM/XSECT

    CALCULATE THE COST OF THE REACTOR

    VS=PI*15.»»2.«H
    SEGM=A1NT ( TOTV/VSI+1.
    1F(H-8.0)38O8.36
 36 CSTR=14.3«30.»H»SEGM
    GO TO 39
 38 CSTR-2000.»SEGM

    CALCULATION OF PRESSURE DROP

 39 CALL PDROPI J.RHO.EMU.GUX.RE.DP)
    CALCULATE COSTS OF POWER AND FAN AND TOTS

    DAP=2.602E<.«DP
    CSTF=85.8»DAP«»0.699
    CSTP=THETA»CKW»KW
    CSTOT-CSTR+CSTF+     CSTP
    WRITEI3t520)XSECT.SEGM,CSTFtCSTP.CSTR,CSTOTi    H   tDP
 <»0 CALL DATSWU.IO
    GO TO 160.50) ,K
 50 0«D+OD
    GO TO 35
 60 GO TO 10
500 FORMAT (!Hlt3X i 'J=' I 2 <2X> • COST /KWH° ' F9.5 >2X ' TOT . VOL. = ' E12.4 »2X
   1 •DIAM.='E12.5.2X1DO»lE12.5.2X'THETA=«E12.5//3Xt1RHO«1E12.S.2X1MU=t
   2E12.5.2X.'MASS VEL.« ' F.12.5 .2X» TF = 'F8.2 . 2X 'MW« ' F8.2 »2X« P= • F 10.4/ / )
510 FORMAT(T6'XSECT'T20'SEGMIT341CSTF'T<»9ICSTPIT62>CSTR1T761CSTOT'T90
   1'  H'TIO^'DELT.P'/I
520 FORMAT(9(2X.E12.5) )
    END
                 • MONSANTO RESEARCH CORPORATION •

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          APPENDIX V
  RELATIONSHIP OF W/F FACTORS
      AND SPACE VELOCITY
'• MONSANTO RESEARCH CORPORATION •

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                             APPENDIX V
          RELATIONSHIP OF W/F FACTORS AND SPACE VELOCITY
   W   =  catalyst weight, g
  dc   =  catalyst density, g/liter
  Lc   =  liters of catalyst
   v   =  total gas flow, cc/sec
   V   =  total gas flow, liters/sec = v x 10~3
  V   =  S02 flow, liters/sec
   N   =  moles total gas/sec
   n   =  moles S02/sec
   P   =  total pressure, atm..
   T   =  temperature, °K
   R   =  gas constant = 0.082 liter atm/mole-°K
                                          V
fS02   =  fraction S02, mole or volume = -
                                           V
   S   = space velocity, sec"1 = liters S02/liter catalyst - sec =
         V
                                            W
Convert W/F, g-sec/cc total gas flow, Q! = -—

to W/F, g-sec/mole S02,               Q2 = -jj
(1)  V = v x 10- 3
(2)  H,-g-
                     f    PV
(3)  n = fSQ2 (N) ,=
                            for out test system:  P = 1 atm
                                                  T = 25°C (298°K)
                                  150
                    • MONSANTO RESEARCH CORPORATION •

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                               (v)
         (0.082)(298)        24.5 x



then
      -\ W(2*J.5 x 103)   _  24.5 x  103 Q,
      ~ —w—ff   \	  ~	
          v  (fso2)
             (PT)
and


(8) S
Convert Q  to S.

For the general case:


                      -"' •  -r™H
but   W = dcLc
            V1

and  fS02 = ~V
          dcLc(RT)       dcLQ(RT)   _     /RT
*+)
                            151




                • MONSANTO RESEARCH CORPORATION •

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                      APPENDIX VI
            COMPARATIVE COST OF PLATINUM

               AND  VANADIA  CATALYSTS
Preceding page blank
                         153
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                              APPENDIX VI

          COMPARATIVE COST OF PLATINUM AND VANADIA  CATALYSTS
The literature fails to reveal catalysts for this application better
than vanadla or platinum, which are about equivalent in conversion
efficiency.  The possibility cannot be ignored that a better cata-
lyst may be announced at any time in the future.  Nevertheless, at
the moment, there are only the two cited.  A cost analysis of platinum
and vanadia catalysts summarized, in Table JJO , indicates that the
platinum catalyst must be at least ten times as effective as vanadia
catalyst to be comparable in cost with vanadium catalyst.  Decrease
in platinum catalyst cost is not expected in the near future because
platinum and labor costs are increasing.
                                  .154


                     • MONSANTO RESEARCH CORPORATION •

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                           Table  40
     SUMMARY OF ECONOMICS FOR VANADIA  AND PLATINUM CATALYSTS(3)
                                V Catalyst         Ft Catalyst
Volume, Cu ft^1)                   30,300               3,020
Initial cost, dollars^2)        1,250,000          12,800,000
Regeneration cost, dollars(2)   	           1,032,000
Capitalized cost, dollars^2)    3,920,000          20,380,000
NOTES:  (1)   Estimate of the volume of platinum catalyst that
              would be equal in cost to the estimated cost of
              vanadia  catalyst.
        (2)   Estimates are based on the same W/F ratio for
              platinum and vanadia  catalysts.
        (3)   Assumptions used for comparison of platinum and
              vanadia  catalysts:
              General:
              a.  rate of return:  8% per year
              Vanadia  Catalyst:
              b.  amount used for the oxidation of S02 in flue gas
                  30,300 ft3
              c.  price:  $4l/ft3
              d.  packed density:  36.8 lb/ft3
              e.  catalyst life:  5 years
              Platinum Catalyst:
              f.  catalyst contains 0.5$ platinum
              g.  packed density:  50 lb/ft3
              h.  catalyst cost:  cost of platinum + $2.75/lb
                  catalyst, (for manufacturing cost)
              i.  regeneration cost:. $0.75/lb catalyst
              J.  regeneration interval:  2 years
              k.  loss of platinum during regeneration:  2%
              1.  price of platinum:  $110/oz (troy)
              m.  catalyst life:  30 years (including obsolescence
                  and/or abandonment)
                                155
                    • MONSANTO RESEARCH CORPORATION •

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 (1)  To estimate the volume of Pt catalyst that would equal
 in cost the cost of vanadla  catalyst required for a particular
 flue gas application.

 Initial cost of vanadia  catalyst
                          =  860,000 liters x ^'^
                                 '             liter

                          =  $1,250,000

Capitalized cost for vanadia  catalyst

                          =  $1,250,000 x
                          =  $1,250,000 x 3-131

                          =  $3,920,000.

Assuming that the amount of Pt catalyst required = P Ib

Initial cost of Pt catalyst,         .
                       A  =  Cost of Pt -t manufacturing
                                  cost of Pt catalyst
                          =  0.005 P Ib x ii7|0  , $2^75  x p lb
                          =  8.8P + 2.75P

                          =  11.55P
                             156
                 •  MONSANTO RESEARCH CORPORATION •

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Regeneration cost of Pt catalyst:
                    H  =  Refining cost + cost of Pt loss
ID
                                 x  P Ib  +  0.0001 x P Ib

                                            $1760
                       =  0.75P + 0.176P


                       =  0.926P


To calculate the volume of Pt catalyst, capitalized cost of

   vanadiu  catalyst  =  capitalized cost of initial.Pt catalyst +

   capitalized cost of regenerating Pt catalyst every two years


            3,920,000  =  AF30 + H (G2-G30)
                       =  11.55P
                          0.926P
       (1+0.08)30
       (1+0.Ob)00
                                  (1+0.08)^-1
                        (1+0.
                       =  11.55P x 1.110 + 0.926P (6.010-0.1103)
                       =  11.55P x 1.110 + 0.926P x 5.8997
                       =  18.283P
                          151,000 Ib of Pt catalyst
                       =  3,020 cu ft of Pt catalyst
                                 157
                     • MONSANTO RESEARCH CORPORATION •

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(2)  at same W/F values of Pt and V catalysts, what weight of

Pt catalyst is required?  Cost of Pt catalyst?  Cost of V cata-

lyst?



    Weight of V catalyst  =  30,300 ft3  x 36^.,lb



                          =  1,115,000 Ib



                          =  505 x 106 gm



    Weight of Pt catalyst =  1,115,000 Ib



                      S02 =  150 x 106 SCFH x 0.003



                          =  0.45 x 106 SCFH



                                    10* — x    Hr    v lb-mole x
                                    xu      x      sec x         x
106  _    ,.     2 mole
  H  '  L'^ x 1U
                             136 x 10H  '   '         sec
                      W/F =  505 x 10^ g x       sec
                                                     mple
                               -37 x 106 g~sec
                              '*< x 1U  mole S02
    Initial cost of V catalyst  = $1,250,000



    Capitalized cost = $3,920,000



    Initial cost of Pt catalyst,
                          =  Cost of Pt + Manufacturing cost of

                                      of Pt catalyst
                                   158
                     • MONSANTO RESEARCH CORPORATION •

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                     =  0.005 x 1,115,000 Ib x
                                   x  1,115,000 Ib
                     =  $12,880,000





Regeneration cost of Pt catalyst





                     =  Refining cost + cost of Pt loss





                     =  ^yp-  x 1,115,000 Ib + 0.0001 x




                             1,115,000 Ib x $1?k°






                     =  $1,032,000





Capitalized cost of Pt catalyst




                     =  $12,880,000 x 1.110 x $1,032,000 x 5.8997





                     = $20,380,000
                                159
                     • MONSANTO RESEARCH CORPORATION •

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                            APPENDIX VII

              SORPTION ISOTHERMS FOR MOLECULAR SIEVES
Molecular sieves have recently been shown to be useful in many ap-
plications as heterogeneous catalysts and as vehicles in studies
of catalysts and catalyst mechanisms.  However, no publication,
to our knowledge, has revealed specific studies to determine their
merit as support material for known S02 oxidation catalysts.

We obtained samples of several experimental forms of zeolite materi-
als and embarked upon a program to screen the materials as potential
supports.  The monodisperse nature of the zeolite pores and their
extremely high specific area are characteristics highly favorable
for maximum contact conditions between the S02 and active catalyst.

Zeolite materials are well known for their adsorption characteristics
with regard to the various constituents in stack gas, and we wished
to know if these characteristics would be an aid or hindrance in our
contemplated use of them as a support material.  Initial experiments
were performed at a low temperature range (75° to 250°F) as shown
in Figures 55 through  6l.
                                162


                   • MONSANTO RESEARCH CORPORATION •

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   901
c
o

a

s  Y0
•a
5  50 J

rt
a
a

<  40
   30-
                                               SK-20

                                               Bed Wt. =6.9 gms
              12    20    28     36    44    52    60    68    76    84


              Exposure Time at 40cc/sec Stack Gas Flow, minutes
       Figure 55.  Sorption for Molecular Sieve  SK-20
                              163
                 • MONSANTO RESEARCH CORPORATION •

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c
o
o
to
T3
 (M

O
CO

c
(b
JH
m
a
a
   100

   lOO
    90-
    80-
    70-
    60-
    50
    40-
    30-
                                               SK-400

                                               Bed Wt. = 6. 7 gms
          4    12    20   28    36    44    52    60    68    76     84

               Exposure Time at 40cc/sec Stack Gas Flow, minutes



         Figure 56.  Sorption Isotherms for Molecular Sieve SK-400
                  • MONSANTO RESEARCH CORPORATION •

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c
0
o
(ft
TD
 (M
O
tfl
a
a
    100
     90
80-
70
60
     50
     40
     30-
                                               SK-^00
                                               Bed Wt.
                                                  = 9.5 gms
          4     12     20    28   36    44    52    60    68     76

                 Exposure Time at 40cc/sec Stack Gas Flow,  minutes
                                                               84
         Figure 57.  Sorption  Isotherms for Molecular Sieve SK-400
                                 165
                   • MONSANTO RESEARCH CORPORATION •

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c
0
4_J
a
L.
0
    90
    HO
70
O   60
to
O.
a.
     30-
                                              SK-^10
                                              Bed Wt. = 6.2 gms
         4     12    20    28   36    44    52    60   ,68     76
                 Exposure Time at 40cc/sec Stack Gas Flow,  minutes
                                                               84
       Figure  58.   Sorption Isotherm for Molecular  Sieve SK-410
                                 166
                    • MONSANTO RESEARCH CORPORATION •

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c
o
o

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c
o
0
V)
O


c
rt
a
a
    90
    80
    70
    60
    50
    40
    30-
13X

Bed Wt. =8.9 gms
         4     1Z   20    28    36    44    52    60    68    76    84

               Exposure Time at 40cc/sec Stack Gas Flow, minutes



          Figure  60.   Sorption  Isotherms for Molecular  Sieve 13X
                                168
                  • MONSANTO RESEARCH CORPORATION •

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c
o
o
(/>
•u
PvJ

O
   90
80'
    70'
   60'
?  50
Q.
a
   40-
   30.
                                             SK-110

                                             Bed Wt. = 6.4 gms
              12    ZO    2«     36     44     52    60    68    76

              Exposure Time at 40cc/sec Stack Gas Flow,  minutes
                                                                84
      Figure  61.   Sorption  Isotherms for Molecular Sieve SK-110
                              169
                 • MONSANTO RESEARCH CORPORATION •

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                       APPENDIX VIII
          TYCO MODIFIED  CHAMBER PROCESS REPORT
Preceding page blank          171



             o MONSANTO RESEARCH CORPORATION •

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                           APPENDIX VIII

         TYCO MODIFIED CHAMBER PROCESS FOR REMOVAL OF SO?

                     FROM POWER PLANT FLUE GAS *
The basic reactions which describe both the Lead Chamber Process
for sulfuric acid manufacture and the basic Tyco Modified Chamber
Process are shown  on  page  175«  Sulfur dioxide is oxidized with
excess nitrogen dioxide and water to form sulfuric acid and equi-
molar quantities of nitric oxide and nitrogen dioxide (1 and 2)
which are absorbed in sulfuric acid forming nitrosylsulfuric acid
(3).  The oxides of nitrogen are recovered by heating the HNS05
after which the NO is reoxidized to N02 CO.  In the Tyco Process,
excess N02 is absorbed in water forming nitric acid (5).

The flow sheet in Figure 62 shows the basic difference between the
standard.Lead Chamber Process and the Tyco Modified Process.  In the
standard process the NO is present in high enough concentrations to
allow the reoxidation to take place in the same reaction mixture as
the S02 oxidation.  In the new process, the NO is too dilute to
allow the slow oxidation to occur in reasonable contact times, thus
requiring the concentration of the NO and separate oxidation.

The process as shown in Figure 62 is profitable, but suffers from
the disadvantage of requiring 31% additional coal to provide the
necessary energy.  The heat is needed to (1) vaporize the NO and
N02 for recycling, and (2) to concentrate the acid back to 80$
for recycle.

In trying to overcome these deficiencies, two major breakthroughs
in process concept have been achieved which may well make the
Modified Chamber Process a practical one for removing S02 from
power plant flue gas and may have a profound effect on the sulfuric
acid manufacturing Industry in general.

Figure 63 shows an Isothermal Scrubber in which the gas is scrubbed
with hot sulfuric acid.  The entering hot acid has an equilibrium
vapor pressure of water such that all the water that comes in with
the raw flue gas leaves with the clean stack gas.  The bottom of
the scrubber is also run at conditions such that there is no net
transfer of water between gas and liquid.

To accelerate the recovery and reoxidation of the oxides of nitrogen
we examined the stepwise reactions  on  page  178.  The third equation
shows that the net reaction is an oxidation of HNS05 which might
well be carried out in the liquid phase.  This reaction does not
occur by simply contacting air with HNS05 solution (in P^SO^), but
we have found that by using an activated carbon catalyst, the
reaction does take place.  Experimental results comparing charcoal-

"Report presented by Tyco at June 11-13* 1969 Contractors meeting.
                                 1,72


                   • MONSANTO RESEARCH CORPORATION •

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catalyzed and uncatalyzed reactions are ;shown on page 179.  Figure 6^
shows the proposed catalytic stripper which takes advantage of the
greatly accelerated N02 recovery.

These two modifications lead to the simplified Catalytic Chamber
Process shown in Figure 65.  The important thing to note is that
no additional heat energy is required above that which comes in
with the flue gas at 300°F.

The sulfuric acid concentration in the recycle (product) stream of
the newly modified system is no longer limited by equilibrium of
vapor pressure considerations as it was in the original Chamber
Process.  It is only limited by the loss of HaSOu vapor at the
temperatures required to permit the water vapor to escape with the
stack gas.  This places a limit of 92% H2SOi+ on the power plant
flue gas cleaning process, but does not impose the same limit on
the Lead Chamber Process which does not have the problem of water
removal.  Acid at 100$ concentration could conceivably be produced
in this simpler Lead Chamber Process using catalytic stripping
techniques.

This removal of the additional heat requirements greatly improves
the economics of the system as shown on pages, 182. andt 183. Use of the
NAPCA guidelines for the catalytic process have raised the total
capital cost and general operating cost estimates to somewhat
higher levels than the more liberal estimates for the baseline
process, but the dramatic difference lies in the saving on heat and
cooling water.   This saving amounts to more than $6 million in
annual operating expenses.  Using credits shown on page.183,  we
estimate the process to earn $2.62 per ton of coal or 39-5? return
on capital investment making the process both technically and
economically attractive.
                               17.3
                   • MONSANTO RESEARCH CORPORATION •

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PROCESS REACTIONS








  (1)        S02 +  2N02   - *•   S03 + NO  +  N02




  (2)        S03 +  H20   - 5-
  (3)        NO  +  N02  + 2H2SOi,  - »-   2HNS05  + H20




  (i|)        2NO + 02    - =-   2N02








  (5)        3N02  + H20    - ^   2HN03 +  NO
                             174





               • MONSANTO RESEARCH CORPORATION •

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BASELINE PROCESS
                      Ho SO,.  80% 80°F



flue gas
• 300°F
2 0.3% SO,
O *
z
en
Z
H
O
rn H- 1
m TO !
JMC
0 VJl tl.
cr\
O po
o
TJ
O

H
O
Z

















REACTOR


to stacK -^ 	


330°F
\







H2° COOLER
I




]•» L*MM —

HN03
ABSORBER
product
HNO








3








N203
• I


NO
OXIDIZER
STRIPPER











stripper
2100°F



SCRUBBER




H2S04 76$
HNS05
275°F




















I







gas to
stack
r








H,SOU
7 O a



clean
gas —
2100°F
1






I




I













f ACID
> POOT FR
* \yWWJ_f£jJl






^ product
H2S04







ACID
CONCENTRATOR







-------
                 ISOTHERMAL SCRUBBER
             stack gas
             250°F
             1% H20
             60 ppm
             N203
flue gas
250°F     	
1% H20
3000 ppm N203
      80*
250°F
.001 M HNS05
H2S0.4 80$
250°F
.11 M HNS05
                    Figure 63
                    176
        • MONSANTO RESEARCH CORPORATION •

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NITROSYLSULFURIC ACID OXIDATION
              i»HNS05  + 2H20   - »-    2N203  + 2H2SOi»




              2N203  + 02   - *•
                       02 + 2H20
                            17.7





               •  MONSANTO RESEARCH CORPORATION •

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      NITROSYLSULFURIC ACID OXIDATION EXPERIMENTS
Carrier Gas


 Nitrogen

 Air

 Air
Packing


Saddles

Saddles

Charcoal
Concentration (mm Hg)
    NO        NO,
   0.01      0.01

   0.01      0.02

   0.01      0.75
                          178
             • MONSANTO RESEARCH CORPORATION •

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          CATALYTIC STRIPPER
                       250°F
                       .11 M HNS05
flue gas
250°F
N02
                              charcoal
                              packing
       flue gas
       250°P
H2SOt, 80$
250°F
.001 M HNS05
                Figure
                 179
    • MONSANTO RESEARCH CORPORATION •

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                 CATALYTIC CHAMBER PROCESS
                   stack gas
                   250°F
                   1% H20
     COOLER
         REACTOR
                              flue
                              gas
                              1% H20
                       COOLER N203
   HN03
   ABSORBER
           \	\
gas
100°F
            product
            HN03 52%
flue gas
300 °P
.3% S02
  COOLER
gas

N02
250°F
flue gas
250°F
                                          ISOTHERMAL
                                          SCRUBBER
                      H2S04
                      .11 M HNS05
                      250°F
                      CATALYTIC
                      STRIPPER
"*"  product
    H2SO,,
                           Figure  65
                              180
                 •  MONSANTO RESEARCH CORPORATION •

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                                            ECONOMIC ANALYSIS
                    ITEM
2
o
2
2
H
O
m
(A
m
>
2)
O
I

O
o
3
TJ
O
oo
Major equipment


Total Capital Cost



Annual Operating Cost


  Heat


  Water



Total


Saving (heat and water)


Other operating costs



Total operating cost
                                            BASELINE PROCESS


                                               $4,535,000


                                                9,2717000
                                                5,220,000


                                                1,133,000
                                                    6,353,000
                                                2,159,^00


                                               $8,513,^00
CATALYTIC PROCESS


   $ 3,985,000


    12,600,000
                                                                $6,185,000
       168,000



       168,000
     3,688,900


   $ 3,856,900
O
z

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                                                      BASELINE PROCESS
                                                       CATALYTIC  PROCESS
o
z
z
H
O

3)
m
(A
m
O
I

O
o
3)
TJ
O
O
Z
    oo
Credits


  H220k 6 $20/ton (100%)


  HN03 % $84/ton



Total credits


Operating costs



Profit


  $/ton coal



Return on Investment
$6,201,000


 3,530,000



 9,731,000


 8,513,^00



 1,217,600


     0.55



    13-1?
$6,201,000


 3,530,000



 9,731,000


 3,856,900



 5,874,100


     2.62



    39.5?

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              APPENDIX IX




VAPOR PRESSURE APPARATUS AND  PROCEDURES
                  183



     • MONSANTO RESEARCH CORPORATION •

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                                                        Glass Tubing
2
en

Z
H
O

3)
m
w
m

a
o
i

o
o
X
TJ
O
O
Z
                    Heater & Controller,
oo
-Cr
                   100ml Flask
                   Sulfuric Acid with
                     Dissolved HNSO,
                                                                                  Rubber Tubing & Pinch Clamps
                                                                               To Vacuum &Liquid N2&Dry Ice Traps'
u
                                                             Magnetic Stirrer
                                                        Figure 66

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                             APPENDIX IX
 PROCEDURE FOR OBTAINING VAPOR PRESSURE DATA

 1.  A quantity of HzSOi* of the desired concentration was weighed
     and transferred into the 100 ml flask (approx. 50-60 ml).

 2.  A quantity of anhydrous nitrosylsulfuric acid was weighed and
     transferred to the flask.
                  Formula:     NzOa *  (803)2
 3.  The flask containing the above mixture was cooled in an ice
     bath to approx. 2°C.

 4.  The flask was connected to the rubber tubing and vacuum was
     pulled on the system (down to approx. 0.05 mm Hg).  This removed
     dissolved gases in the acid as well as the gases from the vapor
     space.

 5.  The pinch clamp above the flask was closed tightly.

 6.  The flask was then placed in the oil bath and heated to the
     desired temperature (while stirring).

 7.  The flask was held at the desired temperature for 1-3 hours.
     (Until no evidence of gas evolution was observed).

 Q.  The remainder of the system was evacuated to approx. 0.01 mm Hg
     and the pinch clamp to the vacuum source was closed.

 9.  The pinch clamp above the flask was opened for about 3 seconds.
     This is long enough to equalize the pressure throughout the
     system but not long enough for a significant quantity of vapor
     to evolve from the solution.

     NOTE:   The stirring bar was stopped prior to opening the pinch
            clamp .

10.  The pressure in the manometer was quickly measured and corrected
     for the ratio of volumes and the temperature difference.  The
     two volumes were (1) the free space above the flask  up to the
     pinch  clamp, and (2) the volume from that pinch clamp to the
     rest of the system including the manometer leg.  The temperature
     of the system was taken as the "weighted" average between the
     temperature of the oil bath and room temperature.

11.  The contribution of water vapor was subtracted from  the total
     corrected pressure.
                                 185

                     • MONSANTO RESEARCH CORPORATION •

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12.  It was assumed that the total pressure (minus H20 pressure)
     was a result of equal moles of NONOa.   No association of these
     gases or any other gases was assumed.


The step which proved to be the cause of the erroneous results was
step 9-  Apparently sufficient vaporization occurs (even with a
quiesent liquid) in two-three seconds to cause a "high" vapor pressure
reading.  This was demonstrated by opening the pinch clamp above the
flask and waiting until the entire system was at equilibrium.  The
final pressure measured at equilibrium differed by about 25% from the
pressure measured after two-three seconds exposure to the flask.
                                 186

                    • MONSANTO RESEARCH CORPORATION •

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          APPENDIX X
SQ3-PRODUCT ACID RELATIONSHIPS
              187



 • MONSANTO RESEARCH CORPORATION •

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00

00
                  HI
                  a
                      500
                     450
                      400
                      350
                      300
                      250
                      200
                                	_l	
                              H	
                               [


                        10
100
1000
                                                SO, IN FLUE GAS, ppm
                                                                                 t
10.000
                   Figure 6?.  Effect of S03  Concentration  in Flue Gas
                               on  Dew Point of  the Acid

-------
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                (D ct
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   •n
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                                 100
                                  90
                                  80
                           70
                                  60
                       Cj  50
                       v>
jjj   40

cc
                                  30
                                  20
                                  10
                                   100     150     200     250    300    350     400

                                                                       DEW  POINT,  °F
                                                                                   450
                                                                                   500
                                                                             550
600
65O

-------