PB 198 408
APPLICABILITY OF REDUCTION TO SULFUR
TECHNIQUES TO THE DEVELOPMENT OF NEW PROCESSES
FOR REMOVING SO2 FROM FLUE GASES FINAL
REPORT - VOLUME II
November 1970
NATIONAL "ECHNICAL 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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INDUSTRIAL CHEMICALS DIVISION
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APPLICABILITY OF REDUCTION TO SULFUR
TECHNIQUES TO THE DEVELOPMENT OF
NEW PROCESSES FOR REMOVING SOg
FROM FLUE GASES
Final Report - Volume II
Phase II of Contract PH-22-68-24
November 1970
For:
New Process Development Section
Division of Process Control Engineering
National Air Pollution Control Administration
U. S. Department of Health, Education, and Welfare
Industrial Chemicals Division
Allied Chemical Corporation
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?H£E|T j e L • i 1 JlfTO-0660 1
t. Title and Subtitle
Applicability of Reduction to Sulfur Technique* to the
Development of Hew Proceaaea for Removing SO2 From Flue Gaaea
II
7. Author(a)
8. Performing Organization Kept.
No.
9. Performing Organization Name and Address
Industrial Chaaicala Division
Allied Chenleal Corporation
P. 0. Box 405
Morristown, New Jersey 07960
10. Project/Task/Work Unit No.
11. Contract/Grant No.
PH-22-Q8-2A
12. Sponsoring Organization Name and Address
Raw Proceaa Davalopcoaat Boot ion
Dlviaion of Proceea Control Engineering
Rational Air Pollution Control Administration
U. 8. Departocnt of Bee Ufa, Education, and Welfare
13. Type of Report & Period
Covered
Final
14.
IS. Supplementary Notes
16. Abatracta
Volume II of a two-volume report which summarizes the results and conclusions of a
study of S02 reduction processes. This Volume covers the Phase II activities of the
study. Four of its five sections deal with experimental studies designed to optimize
process conditions and to confirm the validity of assumptions made in Phase I studies
These sections cover (1) Glaus Process Kinetics, (2) Intermediate Reactor studies,
(3) Low temperature Claus Process studies, and (4) Strong S02 Reduction studies. The
fifth section is a. Phase I type study on the use of dimethylaniline to gather the S02
from smelter gases and deliver a concentrated S02 gas to a reduction process.
17. Key Words and Document Analysis. 17o. Descriptors
S02
Reduction (chemistry)
Flue gases
Kinetics
Mathematical model
17b. Identifiers/Open-Ended Terms
Claus Process
17e. COSATI Field/Group
]3/B
18. Availability Statement
Unlimited
19.. Security Class (This
Report)
UNCLASSIFIED
20. Security Class (This
Page
UNCLASSIFIED
21. "Nfo.'of Pages
155
22. Price
FORM NTIC-S8 (IO-70)
U3COMM-DC 40328-P7 I
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DISCLAIMER
Thia report vaa furnished to the Air Pollution
Control Office by
Industrial Chemicals Division
Allied Chemical Corporation
P. 0. Box 405
Morristown, New Jersey 07960
in fulfillment of Contract No. PH-22-68-24
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FOREWORD
This two-volume report summarizes the results and conclusions
made by the Industrial Chemicals Division of Allied Chemical Corpora-
tion for the National Air Pollution Control Administration under
Contract No. PH-22-68-24.
Volume I covers Phase I activities performed between
June 1, 1968 and July 31, 1969. The objective of Phase I was to
establish the state-of-the-art of reduction to sulfur techniques.
Based on a comprehensive literature survey, thirty case studies
were worked up covering several types of sulfur oxide stack
emissions and several reductants. Each case was based on the
best information available, and use of updated technology in
devising the process sequence. Flow sheets, operating parameters,
and economics are reported in this volume.
Volume II covers Phase II activities performed between
August 1, 1969 and September 30, 1970. Four of its five sections
deal with experimental studies designed to optimize process conditions
and to confirm the validity of assumptions made in the Phase I studies.
These sections cover (1) Claus Process Kinetics, (2) Intermediate
Reactor Studies, (3) Low Temperature Claus Process Studies, and
(4) Strong SOS Reduction Studies. The fifth section is a Phase I
type study on the use of dimethylaniline to gather the S02 from
smelter gases and deliver a concentrated SOa gas to a reduction
process.
The contract work reported herein covers an important
approach to SOa pollution control. It shows the reduction to sulfur
technique to be a viable vehicle. By establishing certain process
applications to be either technically or economically untenable,
it narrows the area to a few preferred routes. An example is the
handling of smelter gas by a gathering process, with subsequent
reduction of the concentrated SOa to sulfur. Laboratory investigations
have optimized operating conditions for selected process steps,
specifically the Intermediate Reactor and the Claus units, that are
common to almost all reduction flow sheets.
Although many departments within the Allied Chemical complex
contributed to this work, Industrial Chemicals Division's R& D staff
held prime responsibility for its planning, direction, and completion.
The following individuals had major time participation in this effort:
Mr. C. A. Bernales, Intermediate Reactor and Strong Gas Studies in
Phase II; Mr. S. B. Boucher, engineering support in Phase I;
Mr. R. L- Burrell, Low Temperature Claus Studies in Phase II;
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FOREWORD
(Continued)
Mr. R. H. Edgecomb, process engineering in Phase I; Mr. G. B. Falk,
economic evaluations in Phases I and II; Mr. T. S. Harrer, engineering
support in Phase I; Mr. R. S. Park, process engineering in Phases I
and II; Dr. L. P. Sharma, general consultant and Strong Gas Studies
in Phase II; Mr. R. L. Sturtevant, general consultant in Phase II;
Dr. S. N. Subbanna, Glaus Kinetic Studies in Phase II; and
Mr. A. W. Yodis, Project Director of Phases I and II.
The cooperation of our NAPCA Project Officer,
Mr. G. L. Huffman, is especially appreciated.
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TABLE OF CONTENTS
1.1
1.2
1.3
1.4
1.5
1.6
1.7
INTRODUCTION
SUMMARY
CONCLUSIONS AND RECOMMENDATIONS
BACKGROUND (PRIOR ART) ___
THEORY
LABORATORY EXPERIMENTAL WORK
1.6.1 Apparatus
1.6.2 Experimental Procedure
1.6.3 Experimental Data and Discussion of Results.
THE KINETIC MODEL _
1.7.1 General
1.7.2 Sub-Systems
1.7.3 The Total System
1-1
1-3
1-4
1-5
1-9
1-12
1-12
1-13
1-14
1-18
1-18
1-21
1-25
1.8 BIBLIOGRAPHY _ __ 1-29
EXHIBIT NO. 1-1
EXHIBIT NO. 1-2
EXHIBIT NO. 1-3
EXHIBIT NO. 1-4
EXHIBIT NO. 1-5
EXHIBIT NO. 1-6
EXHIBIT NO. 1-7
EXHIBIT NO. 1-8
EXHIBIT NO. 1-9
EXHIBIT NO. 1-10
EXHIBIT NOo 1-11
EXHIBIT NO. 1-12
EXHIBIT NO. 1-13
EXHIBIT NO. 1-14
EXHIBIT NO. 1-15
EXHIBIT NO. 1-16
EXHIBIT NO. 1-17
EXHIBIT NO. 1-18
Experimental
Experimental
Temperature
H20)
Temperature
Experimental
Experimental
Experimental
Experimental
Experimental
Experimental
Experimental
Experimental
Experimental
Experimental
Experimental
Experimental
Reaction _.
Experimental
Experimental
H20 + C02)
Data (H2S + S02 + H20).
Apparatus
Profile (COS + S02 +
Profile~~(H2S~+~s62)~"I'
Data (H2S + S02)
Data (H2S + S02)
Data (H2S + S02)
Data (H2S + S02 + H20).
Data (H2 + S02 + H20)_.
Data (COS + N2)
Data (COS + S02)
Data (COS + H20)
Data (COS + S02 + H20).
Data (CO + S02 + H20)_.
Data (CO + H20)
Data of Overall
Data (H2S + S02 + CO)..
Data (H2S + S02 +
1-31
1-32
1-33
1-34
1-35
1-36
1-37
1-38
1-39
1-40
1-41
1-42
1-43
1-44
1-45
1-46
1-47
1-48
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TABLE OF CONTENTS
(Continued)
EXHIBIT NO. 1-19
EXHIBIT NO. 1-20
EXHIBIT NO. 1-21
EXHIBIT NO. 1-22
EXHIBIT NO. 1-23
EXHIBIT NO. 1-24
EXHIBIT NO. 1-25
EXHIBIT NO. 1-26
EXHIBIT NO. 1-2?
EXHIBIT NO. 1-28
EXHIBIT NO. 1-29
EXHIBIT NO. 1-30
EXHIBIT NO. 1-31
Equilibrium Constant 1-49
Analog Schematic (H2S + S02) -- 1-50
Arrhenius Plot (H2S + S02) - 1-51
Analog Schematic (COS + S02) 1-52
Arrhenius Plot (COS + S02) 1-53
Analog Schematic (COS + H20) 1-54
Arrhenius Plot (COS + H20) .._ 1-55
Analog Schematic (COS + S02 + H20)_. 1-56
Analog Schematic (CO + S02 + H20) 1-57
Arrhenius Plot (CO + S02) 1-58
Comparison of Results 1-59
Digital Program 1-60
Computer Printouts 1-62
1* •
-11
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1. CLAUS PROCESS KINETICS
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1.1 INTRODUCTION
1.1.1 In recent years the emission of sulfur dioxide into
the atmosphere has increased appreciably and is
posing a serious air pollution problem. Reduction of SOa
to elemental sulfur is one useful approach to the abatement
of sulfur dioxide atmospheric pollution.
1.1.2 Since June 1968, Allied Chemical Corporation has
contracted with National Air Pollution Control
Administration under Contract No0 PH 22-68-24 to study the
applicability of reduction to sulfur techniques to the
development of new processes for removing S02 from flue
gases. The whole program was divided into three phases:
Phase I - A comprehensive literature survey to select
and evaluate potentially useful techniques.
Phase II - Laboratory experimental work to generate
knowledge or data not available in the
literature, to test the assumptions on which
Phase I evaluations rest and to develop or
optimize selected reduction processes.
Phase III - An engineering and economic study of processes
surviving Phases I and II which is sufficiently
detailed to allow a decision on further
development leading to commercialization.
This report deals with the conventional (Normal) Glaus and
is a part of the Phase II effort.
1.1.3 The Glaus process was developed in Germany in 1890.
It consists of the vapor phase catalytic oxidation
of hydrogen sulfide to elemental sulfur by reaction with
sulfur dioxide.
2 H2S + S02 > ^ Sx + 2 H20
The reaction takes place over an alumina catalyst at
temperatures in the range of 392 to 752°F. Although the
Glaus process has been used commercially for many years,
the literature provides no reliable kinetic data to
permit optimum Glaus reactor design. Reliable kinetic
data are also needed to increase Glaus reaction efficiency.
1-1
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1.1.4 Substantially all S02 reduction processes reported
in the literature use the Glaus process in the final
stage of sulfur recovery. Therefore the efficiency of
Glaus operation has some effect on the overall sulfur
recovery. The Glaus process may be considered as the
primary tool for reducing S02 emissions to the atmosphere.
In most cases the exhaust gases from the Glaus unit are
discharged into the atmosphere creating S02 pollution
problem. Therefore an increase in Glaus reaction efficiency
has the beneficial effect of reducing the S02 atmospheric
pollution.
1.1.5 Depending on the upstream reduction steps, the Glaus
feed may contain any or all of the following: S2,
CO, COS, CS2, H2, C02 and H20 in addition to the primary
reactants H2S and S02« How these components behave under
conditions optimal for H2S/S02 reaction can have a marked
effect on sulfur yield. The question arises whether or
not the CO, COS, CS2 and H2 react fully in the Glaus unit
in the residence time provided. If these reactions with
S02 are slower than the H2S reaction, then either a large
catalyst charge must be provided in the unit or loss of
reductant be accepted. Either of these alternates would
involve an economic penalty. It is not possible to answer
any of these questions from the previously existing kinetic
data. Incomplete conversion of CO, COS, CS2 and H2 in the
Glaus unit means incomplete utilization of the full
reducing power of the upstream reductant. This results in
increased cost of the upstream reductant. Depending on
the composition of CO, COS, CS2 and H2 in the feed gases
and the economics of the overall situation, it is necessary
to decide whether or not an Intermediate Reactor is needed.
The main purpose of the Intermediate Reactor is to completely
react the components such as CO, COS, CS2 and H2 with S02
and to ensure that H2S, S02 and S2 are the only sulfur
bearing species entering the Glaus unit. Again no decision
on the Intermediate Reactor can be taken with the available
kinetic data.
1.1.6 It was therefore the aim of this program to develop
a mathematical model for the kinetics of the Claus
process. A further objective of this study was to determine
the reactivity of CO, COS and H2 in the feed gases.
Depending on the reactivity of CO, COS and H2 in the Claus
unit, it was also the aim of this study to provide data
which would allow decision as to whether or not an
Intermediate reactor is needed. It was not the aim of
this program to go into details of the mechanism of
reaction which is a much more complicated problem than the
development of a satisfactory rate equation.
1-2
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1.2 SUMMARY
1.2.1 This report covers the development of a mathematical
model for the kinetics of the normal temperature
Glaus process.
1.2.2 Experiments were carried out for generating the
experimental kinetic data. These data were collected
using the dynamic flow system over Porocel LPD catalyst.
The temperature range was 400 to 700°F and the contact times
were varied from 1/16 to 2.5 seconds. The H2S concentrations
in the feed were varied from 1 to 6 percent. The experi-
mental kinetic data for the Glaus reactions are presented
in Exhibits 1-5 through 1-18.
1.2.3 It was found that the reaction between H2S and S02
is very fast and reaches almost equilibrium conditions
in 1/2 second contact time. The kinetics and the rate of
(H2S + S02) reaction is not affected by 20% water in the
feed. For each temperature in the Glaus range, there is
an upper limit to the inlet compositions of H2S and S02,
above which, liquid sulfur condenses and deactivates the
catalyst. Hence the Glaus reactor must be operated above
the dew point of sulfur. In general it may be said that
(COS + H20) and (COS + S02) reactions are very slow compared
to the (H2S + S02) reaction. A substantial part of COS does
react. Carbon monoxide reacts only to a slight extent even
at 700°F. Hydrogen does not react at all and passes through
like an inert. It is therefore concluded that in order to
completely react the CO, COS and H2 in the feed, it is
necessary to employ an Intermediate Reactor.
1.2.4 The integral method was employed for analysis of
kinetic data. The equations for each of the
sub-systems were written and analog programs used to
evaluate the rate constants. The overall model was built
up stepwise with an effort made to use the least number
of reactions. The final model was built using 8 reactions.
The overall model is a digital program written in Fortran IV.
It includes the complex equilibrium program to make certain
that equilibrium conditions are ultimately reached at
infinite contact time. For a given feed composition and
temperature, the program goes through the complex equilibrium
calculations as well as the Runge-Kutta numerical integration
and prints out the product compositions as a function of
contact time. The listing of the main program and computer
printouts are given in Exhibits 1-30 and 1-31 respectively.
1-3
-------
1.2.5 The model predicts the product distribution reliably
over the normal Glaus temperature range. At higher
temperatures of 700°F, the conversions predicted by the
model are less than those obtained by experiment. The
experimental results at 700°F are considered unrealistic
since the experimental conversions are even better than
what the equilibrium permits. More than equilibrium
conversions in these experiments are explained on the
basis that further reaction must be taking place at lower
temperatures (where equilibrium is more favorable) in the
bottom leg of the reactor. Such a situation does not exist
for lower temperature runs since the catalyst temperature
and the bottom leg temperature are approximately the
same.
1.2.6 The results predicted by the model are conditioned
by the equilibrium considerations and therefore are
quite safe for design purposes. Although the model is built
using 1 to 6% H2S, it is expected on the basis of a few
runs made, to hold good for at least 10% H2S in the feed.
1.2.7 It was found that C02/H2S ratio has no effect on the
conversion of H2S for C02 concentrations of 5 to 20%
in the feed gases.
1.3 CONCLUSIONS AND RECOMMENDATIONS
1.3.1 It was found that the (H2S + S02) reaction is very
fast and reaches almost equilibrium conditions in
1/2 second contact time. Although some runs have been made
in this work at very short contact times of 1/16, 1/8 and
I/A seconds, further more detailed experimental work would
be needed to define exactly the kinetics of this reaction
at very short contact times. These experiments should be
performed in a 1" reactor tube rather than in a 2" reactor
tube used in this work. Using these more accurate kinetic
data, the model can be improved and refined further.
1.3.2 The kinetics and the rate of (H2S + S02) reaction
was found to be unaffected by 20% water in the feed.
The (COS + S02), (COS + H20) and (CO + S02) reactions are
very slow compared to the (H2S + S02) reaction.
1.3.3 It is concluded that in order to completely react
CO, COS and H2, it is necessary to employ an
Intermediate Reactor. The Intermediate Reactor may be
recommended for any particular reduction scheme, depending
on the composition COS, CO and H2 in the feed, the overall
economics and the atmospheric pollution problem.
1-4
-------
1.3.A All work in this study was done by employing
Porocel LPD catalyst. The conclusions drawn are
therefore true only for this catalyst. A catalyst develop-
ment program may possibly uncover another catalyst that
would enhance the reactivity of less reactive constituents.
1.3.5 The model predicts the product compositions reliably
over the Glaus temperature range. The conversions
predicted by the model are less than those obtained by the
experiment around 700°F. However, the experimental results
are considered unrealistic at these higher temperatures.
The use of the model is quite safe for design purposes
even at 700°F. The model is conditioned by the equilibrium
considerations.
1.3.6 It is recommended that more accurate kinetic data
be collected at 700°F before attempting to improve
the model. Improvements in the experimental apparatus and
sampling techniques should be incorporated for avoiding
the further reaction in the bottom leg of reactor. The
exit samples should be taken immediately after the catalyst
bed to eliminate any chance for further reaction.
1.3.7 The model is built up using 1 to 6% H2S in the feed.
It is expected, on the basis of the few runs made,
to hold good for 10% H2S in the feed. Since the model is
semi-empirical, the extrapolation beyond the range of
variables tested may or may not be good.
1.3.8 The model is built considering only SQ as the
elemental sulfur species. This is a good approxima-
tion in the Glaus temperature range. However, at 700°F
other forms of sulfur species such as S2, S< and 85 are
also present to some extent. The model could be improved
by considering these species although the model would
become complicated. No such attempt was made in this
study.
1.3.9 From the point of view of kinetics, 1/2 second
contact time for (H2S + S02) reaction is sufficient.
However the contact time to be used in any real situation
also depends on the overall optimal operation of the plant
and the S02 atmospheric pollution considerations.
1.4 BACKGROUND (PRIOR ART)
1.4.1 The demonstration of the rapid reaction between
hydrogen sulfide and sulfur dioxide in the presence
of moisture was a common lecture experiment as early as
1812, but is is not known by whom this reaction was first
1-5
-------
observed. However Chuzel (1) in 1812 noted that no reaction
would result between hydrogen sulfide and sulfur dioxide if
the gases were first dried by passing over calcium chloride.
1.4.2 Matthews (2) in 1926 confirmed that dry gaseous H..S
and SO- do not react on mixing but they will react
in the presence of a liquid film of water on exposed
surfaces. He did not carry out any quantitative experiments
on H.-:S/S02 reaction.
1.4.3 Lidov (3) in 1928 noted that the reaction between H.^S
and SOc is very rapid and complete. He even suggested
that H2S may be determined by adding a known amount of S02 .
1.4.4 Randall and Bichowsky (4) in 1918 showed that the
rate of reaction between SOo and H2S depends on the
surface of the reaction vessel. They observed that a very
small amount of sulfur is formed when a mixture of moist
S02 and H2S is passed through a clean glass tube at 60°C,
but if the tube is first etched, then the sulfur is formed
at a more rapid rate.
1.4.5 Taylor and Wesley (1) in 1926 studied the kinetics of
the gaseous reaction between ft^S and SO.?. They used
the dynamic flow method and the reactions were conducted in
glass tubes. The gases after reaction were cooled rapidly
to 100°C, where the reaction rate is exceedingly slow.
The unreacted H;jS and SO^ were analyzed by absorbing
them in NaOH solutions. They covered a temperature range
of 371 to 733°C in two reaction tubes of the same volume
but with different surface areas (317 and 121 cm^) . They
found that the reaction rate is determined by the equation
. _ ,
S ~ k
PS0
where S is the amount of sulfur formed in gms/min.
k is the rate constant.
o and pqn are the average partial pressures
o bu2 of R2g and S02 respectively.
They finally concluded that the reaction rate between H2S
and S02 is proportional to the surface area of the reaction
vessel. The reaction takes place almost exclusively on the
surface of the reaction vessel and very little in the gaseous
phase .
1-6
-------
1.4.6 Yushkevich etal (5) in 1932 made equilibrium calcula-
tions for H2S and S02 reaction for the temperature
range of 200 to 800°C and pressure range of 5 to 760 mm Hg,
on the basis of the following reactions:
2 S02 + 4 H2S <—** 4 H20 + 3 $2 (2)
S6 (3)
3 S8 (4)
2 H2S (5)
However the reaction (5) may be ignored for equilibrium
calculations below 700°C because of the formation of
negligible amounts of H2. Above 900°C where Sa molecules
predominate, the reaction (2) is endothermic. Therefore
from reaction (2) alone, one would expect that the higher
the temperature, the more complete the reaction between
S02 and H2S. If reactions (2), (3), (4), and (5) are
considered simultaneously, the equilibrium conditions
change markedly.
1.4.7 Yushkevich etal (5) found that the H2S and S02
reaction is very slow in the absence of a catalyst.
They used catalysts such as Tikhvinsk bauxite
(Si02:Fe203:Al203::8.5:20.4:61.0}, Alapaevsk bauxite
(Si02:Fe203:Al203::4.5:26.5:27.2) and Alapaevsk iron ore
(Fe203:Al203:Si02:: 80.0:14.1:4.0). While using bauxite
catalysts they observed that at 100°C and SVH of 120-130,
the reaction practically goes to completion (98-9970) and
both types of bauxites are equally good. However their
experiments indicate a decrease in the degree of reaction
for dilute gases (H2S:S02::1.5:0.75) and rapid decrease in
reaction rate for temperatures below 100°C. They also found
that at 200 °C and SVH of 1000, the reaction was completed
to 86% with bauxitic iron ore catalyst while Tikhvinsk
bauxite gave only 247» conversion of S02 and H20
(H2S:S02:N2::1.5:0.75:97.75) at the same temperature and
SVH of 725,, They finally concluded that the bauxite iron
ore catalyst is better than the bauxite catalysts.
1.4.8 Lepsoe (6) in 1940 studied the kinetics of S02
reduction by carbon, CO and COS. He found that the
reduction of sulfur dioxide by carbon may be expressed
satisfactorily by the following consecutive reactions
S02 + C > C02 + 1/2 S2
C02 + C > 2 CO
1-7
-------
The- rate of formal ion of CO. between 900 °C and 1200 °C,
expressed as moles in the reaction products, is given by
the formula
(co.,) = i.ii[(so.?) - (so.-?)0'1]
He observed that above 1200°C, the rate of sulfur dioxide
reduction appears to be controlled by gas diffusion rates
and substantially the same depth of fuel bed being
required for the reduction of SO?, regardless of gas
velocity. He also found that the reduction of sulfur
dioxide by means of CO or COS is very fast with any
kind of catalyzing surface above 800°C. At lower tempera-
tures (250 to 500°C) alumina is an efficient catalyst and
the reaction appears to be of the first order. His
experiments show that the reduction with COS is approximately
four times as fast as with carbon monoxide. He used SVH of
60 and reports reaction efficiency of 80 to 100% for COS/S02.
1.4.9 Gamson and Elkins (7) in 1953 presented a review of
literature on the conversion of H2S to elemental sulfur.
In this article, the thermodynamics of the formation of sulfur
from hydrogen sulfide is developed and a rigorous and unique
calculation procedure is outlined. They have made equilibrium
calculations for H2S/Os and H2S/S02 reactions. They give kinetic
and yield data for the reaction of H2S and S02 covering superficial
space velocities of 240 to 1920 SVH, and temperatures from 230 to
300°C. They give similar data for the reaction of COS and S02,
using an SVH fixed at 200 and varying the temperature from 222 to
303°C. Their results show that for a feed gas of 5.5% COS, 2.75%.
S02, and 91.75% N2, the reaction is 97.2% complete at a tempera-
ture of 303°C. They used the dynamic flow method and 4-8 mesh
Porocel catalyst.
1.4.10 For the COS-S02 reaction, superficial space velocity (SVH)
used by Lepsoe is 60 ,and that used by Gamson and Elkins
is 200. Both of these are low and do not cover the full range
of typical normal Glaus operation.
1.4.11 Munro and Masdin (8) in 1967 studied the desulfurizing
of fuel gases. Their equilibrium calculations for H2S
and S02 reaction were in agreement with those of Gamson and Elkins
discussed before. They used 13X molecular sieve as catalyst in
their experimentation. They found good agreement between
experimental results and theoretical predictions. They therefore
concluded that the theory can be used to predice accurately
the conversion when using a fully active catalyst.
1.4.12 On the basis of the existing kinetic data, one cannot
accurately predict the behavior of COS in the Glaus
regime under the usual operating conditions, although it is
clearly evident that H2S is more reactive than COS. Kinetic
data for COS reactivity has been developed covering a wider
range of parameters.
1-8
-------
1.4.13 The discussion so far was an attempt to survey the
existing literature on the kinetics of Glaus process,
It is by no means complete. However, one must conclude that
although the Glaus process has been used commercially for
many years the literature provides no reliable kinetic data
to permit optimum Glaus reactor design.
1.5 THEORY
1.5.1 The main objective of this program was to find an
adequate rate equation for the kinetics of the
Glaus reaction. At present chemical kinetics is not,
however, an exact science. From a practical standpoint
it is not yet possible to formulate generalized mathematical
relations for the rate of a chemical reaction. It is,
therefore, necessary to determine the rate of reaction
experimentally.
1.5.2 Kinetic data for a catalytic system are best obtained
in flow reactors and the method most often used is
the integral reactor. Integral reactors have a distinct
advantage over differential reactors because the chemical
analysis need not be very rigorous for a reasonable degree
of accuracy. The compositions are measured as a function of
feed rate and temperature. Kinetic data obtained in this
manner are the most dependable and simple to use. The
method has the advantage of direct applicability to
flow-type reactors.
1.5.3 The Glaus reactions are heterogeneous vapor phase
catalytic reactions. It is assumed that the reaction
proceeds at all the gas-solid interfaces both at the outside
boundaries and within the porous catalyst pellets. For such
a reaction we select, as the most reasonable representation
of reality, a continuous-reaction model which pictures
reaction occuring to a lesser or greater extent throughout
the catalyst pellets. This is in contrast to the shrinking
unreacted-core model with frts definite zone of reaction
which most reasonably represents the real case in the
majority of non-catalyzed gas-solid reactions.
1.5.4 In developing rate expressions for the continuous
reaction model, the various processes that may cause
resistance to reaction must be taken into account. These
are:
(a) Gas Film Resistance - Reactants must diffuse from
the main body of the gas to the exterior surface of the
catalyst.
1-9
-------
(b) Pore Diffusion Resistance - Substantially all of
the surface area of the catalyst is inside the pores.
Therefore the reactants must in general move into the pellet
through the pores.
(c) Surface Phenomenon Resistance - The reactants
are adsorbed to the surface of the catalyst where they
react to give products. The products are then desorbed back
to the gas phase within the pore.
(d) Pore Diffusion Resistance for Products - Products
then diffuse out of the pellet.
(e) Gas Film Resistance for Products - Products then
move from the mouth of pores to the main gas stream.
1.5.5 Since the steps listed above take place in series,
it is possible for any one of them to control the
overall rate of reaction. The slowest step is known as
the rate controlling step. The experimental data were
collected to obtain a reliable rate equation for Glaus
reactions. Therefore, proper precautions were taken to
make certain the step (c) is the rate controlling step.
1.5.6 The effects of diffusional resistance are kept to a
minimum by using a high velocity through the catalyst
bed. These effects may be tested for in the experimental
reactor by varying the weight of catalyst and the feed rate
(W/F) at the same time (9). The conversion is measured at
a value of W/F at which the gas velocity is low. Then it
is measured again at a high velocity, but with more catalyst
to keep the ratio W/F constant. The values of conversion will
coincide if the effects of diffusion is negligible. If the
conversions are different, there is a diffusional effect.
The experimental kinetic data and the conversions, obtained at
different linear velocities, but at same W/F values are shown
in Exhibit No. 1-1. It may be seen that the conversions are
approximately the same within experimental error.
1.5.7 The following paragraphs discuss rate of reactions.
In contrast to the batch case, consider the system
where the reactants flow continuously into the reactor and
the products are continuously removed. The flow reactor may
be represented by the figure given below.
AZ
CA
1-10
-------
CA is the concentration of substance A in moles/ft3 and
v is the volumetric flow rate in ft3/hour. The subscripts
o and f indicate the entrance and outlet conditions.
The material balance for a reaction component A may be made
for a differential element of length AZ.
Input-Output-Disappearance by Reaction = Accumulation
(vCA)|z - (vCA)|z+Az - (kCA)SAZ = j£ (SAZCA) (1)
It is assumed that the heterogeneous catalytic reaction
takes place with apparent first order with respect to
component A for purposes of illustration. In the above
equation S is the void cross-section of the tube and k is
the reaction rate constant. The partial differential
equation then becomes
_ d fvCi} - k-r»-ljAiJ O -, i_ ——^—^—— {£. J
At steady state
acA
= 0
at
1.5.8 In the Glaus reactions studied in this program, about
90 percent of flow rate is due to inert nitrogen gas.
The change in the volumetric flow rate due to change of moles
by reaction is negligible. Therefore v may be considered as
constant.
-kcAs - (4)
The contact time t is defined by
where Vr is the void volume of the reactor.
From equations (4) and (5), it follows that
u t = ^CA (6)
Under the conditions and assumptions described above, equation
(6) gives the desired rate expression.
1-11
-------
1.5.9 In the past chemical engineers have frequently used
the Langmuir-Hinshelwood approach, which provides the
adsorption terms for correlating the kinetic data for
heterogeneous reactions. This approach does not have the
theoretical validity commonly attributed to it and its use
leads to unnecessary mathematical complexity. It seems
reasonable that the simplest possible rate equation which will
adequately fit the experimental data b£ employed. The following
expression for the rate equations is among the simplest forms
having sufficient generality (10).
rate = k(pA)m(pB)n(pc)0 --- (7)
Where A and B might be reactants and C a product or a foreign
gas. Depending on the accuracy of the experimental data,
the exponents m, n and o may or may not be restricted to
integral or half-intergral values. The exponents in the
above equation may be considered as simply the apparent orders
of the reaction with respect to the individual components.
1. 6 LABORATORY EXPERIMENTAL WORK
1.6.1 Apparatus
1.6.1.1 The main aim of the experimental work was to
generate the kinetic data needed for developing
the rate equation. A schematic diagram of the experimental
apparatus used is shown in Exhibit No. 1-2. It consists of
a 2" reactor where the reactions are carried out with
necessary auxiliary equipment.
1.6.1.2 The reaction gases were metered using
Fischer-Porter rotameters and mixed with
metered gaseous N- so as to obtain the required reactant
gas concentrations. Where desired H.-.Q was introduced into
the gas stream by bubbling the N;? flow through a
thermostated water saturator. The reaction gases were
manifolded and passed downward through the vertical 2" I.e.
Vycor reactor tube containing 10" of the catalyst bed.
The reaction temperatures were maintained by a Lindberg
Heviduty 3-Zone tube furnace. All three zones were equipped
with separate Pyrovane temperature controllers. The controllers
were adjusted to obtain essentially a flat temperature
profile. Typical temperature profiles under actual run
conditions are shown in Exhibits 1-3 and 1-4.
1.6.1.3 After leaving the reactor, the exit gases
were passed through the bottom leg of the
reactor (Pyrex U-tube), which was maintained around 375°F
to prevent sulfur blockage. Before the reaction products
1-12
-------
were vented to the exhaust hood system, they were sampled
and analyzed by gas chromatographic techniques for S02, H2S,
CO, COS, CS2, C02, N2, etc., as appropriate for the system
under study. Provisions were also incorporated for sampling
of the inlet gas stream and, where appropriate, such samples
were taken and similarly analyzed using the gas chromatographic
techniques.
1.6.2 Experimental Procedure
1.6.2.1 The required depth of the catalyst bed was
added to the reactor tube. Reactant gases
of the required composition were passed through the reactor
by adjusting the rotameters of the various gases. Temperature
controllers were turned on to bring the reactor to the
required temperature. In the beginning the water saturator
was by-passed and dry N2 was mixed with other reactant gases.
Two feed samples were taken without using any drying tube
since the gases were already dry. After taking feed samples,
if required, N2 was passed through the water saturator. The
temperature of water saturator was adjusted to give the
percent water required in the reactant gases.
1.6.2.2 The temperatures at various heights in the
catalyst bed were measured and the tempera-
ture controllers were adjusted to get essentially a flat
temperature profile across the height of the catalyst bed.
When N2 was passed through the water saturator, the line
from the saturator to reactor was heated and maintained at
a temperature slightly higher than the saturation temperature.
This was necessary to avoid any water condensation from the
saturated N2 stream. Heat was also supplied to the bottom
leg where the temperature was maintained around 375°F to
prevent any sulfur blockage.
1.6.2.3 The whole system was then allowed to reach
steady state conditions. When everything
was steady the exit samples were taken. The exit sampling
device consisted of a calcium chloride drying tube and
the sample bomb. The exit gas was first passed through the
calcium chloride drying tube to remove sulfur and moisture
and the dry sample was then collected in the sample bomb.
For each run four exit samples were taken. In between
samples, the temperature profile was measured to make certain
that the experimental conditions had not changed.
1.6.2.4 After taking the exit samples, the water
saturator was again by-passed and the dry N2
stream was mixed with dry reactant gases. Two more feed
samples were taken. This was to check that gases had been
flowing at steady rate without any fluctuations. The feed
samples as well as exit samples were analyzed using two
1-13
-------
column gas chromatography. In most cases, the samples were
analyzed within one hour of collection. The samples containing
air were discarded.
1.6.3 Experimental Data and Discussion of Results
1.6.3.1 Exhibits 1-5 through 1-18 present the
experimental kinetic data as they were
obtained from the G.C. analyses. These data were used
during subsequent computer evaluations resulting in the
determination of the reaction rate constants.
Depending on the particular system under study, experiments
were performed from 400 to 700°F. Reactant gas concentra-
tions were varied from 1 to 6 percent. The contact times
were varied from 1/16 to 2.5 seconds. Contact times were
calculated ba.-sed on the reactor temperature and on the
assumption of. 5070 voids in the catalyst bed. It may be
given by
. _ Void volume of catalyst bed in cu. ft.
Contact Time - piow rate of feed gases at the reactor
temperature and atmospheric pressure
in cu. ft./sec.
1.6.3.2 Experience has shown that while rotameters
are accurate to within i"1070, they are also
subject to variations resulting from the slight changes in
back pressure encountered in this work. The variations thus
found in the inlet gas compositions were naturally reflected
to some degree in similar variations in the exit gas analyses.
Since all gas analyses were carried out using gas chromatographic
techniques, they were subject to the errors inherent in this
method. Since the G.C. nitrogen analysis was not reliable,
it was impossible to obtain an accurate check on a 100%
summation basis. Analytical methods did not lend themselves
to direct determination of H20 or sulfur. Where these
constituents are present in the gases exiting the reactor,
their concentrations could only be calculated by differential
mass balance over the system. Little error in this technique
was expected, since the inlet H20 constant was based on the
assumption that saturation was obtained at the controlled
temperature of the water. This method of H2Q addition has been
verified to be quite accurate.
1.6.3.3 All factors considered it is estimated that the
experimental data are within +10-1570 of the
true values. This is further born out by comparison of the
results of duplicate tests, run at different times. Analog
evaluations of the experimental data further substantiated
this conclusion, particularly when weighted consideration was
given to values checked by duplicate runs.
1-14
-------
H2S
1.6.3.4 The experimental kinetic data for (H2S + S02)
and (H2S + S02 + H20) are given in
Exhibit 1-5 through 1-8. In general it was found that the
(H2S + S02) reaction is very fast and reaches almost
equilibrium values in 1/2 second contact time. It was also
found that the kinetics of (H2S + S02) reaction is not
affected by the presence of 20% water in the feed. It was
observed that at 400 and 450 °F, if the feed compositions of
H2S and 80s a*e high, liquid sulfur condenses and deactivates
the catalyst o A feed composition of 1.5% S02 and 3% H2S
was used at 400 and 450°F. Even in these cases some liquid
sulfur condensation were observed on the walls of the reactor
but not on the catalyst. The following reaction was assumed
for the kinetic analyses.
2 H2S + S02 - > 3/8 S8 + 2 H20
It was assumed that mainly 83 is formed in the above reaction
although some amounts of 84 and 85 are also simultaneously
formed.
COS + S02 + H20
1.6.3.5 The experimental kinetic data for (COS + S02),
(COS + H20) and (COS + S02 + H20) are given
in Exhibits 1-10 through 1-13 respectively. The following
reactions are possible.
2 COS + S02 > 3/8 S8 + 2CO£ (1)
2 COS + 2 H20 > 2 H2S + 2 C02
2 H2S + S02' > 3/8 SB + 2 HgO (2)
2 COS + S02 > 3/8 83 + 2 C02
2 COS > CS2 + C02
CS2 + S02 > 3/8 S8 + C02 (3)
2 COS + S02 > 3/8 S +2 C02
2 COS > CS2 + C02
CS2 + 2 H20 > 2 H2S -H C02
2 H2S + S02 > 3/8 S8 + 2 H20
2 COS + S02 > 3/8 S8 + 2 C02
1-15
-------
2 COS > 2 CO + 2/8 S8
2 CO + S02 > 1/8 Ss + 2 C02 (5)
2 COS + S02 > 3/8 Sg + 2 C02
2 COS > 2 CO + 2/8 S8
2 CO + 2 H^O > 2 C02 + 2 H2
2 H2 + SO- > 1/8 S8 + 2 H20
2 COS + SO- > 3/8 Sg + 2 C02
1.6.3.6 An examination of the kinetic data indicates
that a substantial part of COS does react
in the normal Glaus and no CO is formed in the products.
Hence there is no need to account for CO in the rate
equation. Therefore possibilities (5) and (6) are
eliminated. The experimental data also indicate that no
CS2 is formed. Even if COS is decomposing to form CS2 and
C02, the CS2 formed must be reacting completely leaving
no CS2 in the products. Therefore alternates (3) and (4)
are the same as (1) and (2). For these reasons, the kinetics
were developed based on the following reactions.
2 COS + S02 < > 3/8 S8 + 2 C02
COS + H20 < > H2S + C02
2 H2S + SOo < > 3/8 S8 + 2 H20
CO + S02 + H20
1.6.3.7 The experimental kinetic data for (CO + S02 +
H20) and (CO + H20) are reported in
Exhibits 1-14 and 1-15 respectively. These data indicate
that there is no reaction of carbon monoxide in the normal
Glaus range. However, there is some reaction at 600 and
700°F.
Carbon monoxide can react with S02 to form COS and C02.
The COS formed can react in many possible ways as discussed
before. After considering these as well as other possibilities:
it was found that the kinetics may be developed by the
following reactions.
3 CO + S02
2 COS + S02
COS + H20
2 H2S + S02
1-16
^
^ —
' COS + t- C02
— > 7/P. <^n -t- ? rc\
^ TOo + H-S
— > 3/8 So + 2 HoO
(/)
/Q\
^o;
(9)
V. J /
CIO
V ••-"
-------
H2 + SOg + HgO
1.6.3.8 The experimental kinetic data are reported
in Exhibit 1-9. Examination of the kinetic
data indicates that there is hardly any reaction in the
whole range of 400 to 700°F.
Overall Reaction
1.6.3.9 In order to get an overall picture of the
Glaus process, experimental kinetic data on
the overall reaction were obtained. Typical Claus feed
conditions were used in these runs. Necessary care was
taken to see that sufficient S02 was available to react
completely H2S, CO, COS and H2. The experimental kinetic
data are reported in Exhibit 1-16.
1.6.3.10 The results seem to be in order. These
were expected on the basis of the individual
reactions. It is expected that the overall kinetic model
may be developed on the basis of reactions (7) through (10),
given before. These data were later used to check the
overall kinetic model.
Condition of Catalyst after Run
1.6.3.11 The condition of catalyst was found good after
(H2S + S02), (H2S + S02 + H20), (H2 + S02 +
H20) and the overall reaction runs. But after (COS + S02),
(COS + H20), (COS + S02 + H20), (CO + H20) and (CO + S02 + H20)
runs, it was found that catalyst had turned slightly dark.
It is attributed that it was due to carbon deposition on the
catalyst. However, no deactivation of the catalyst was observed,
No catalyst life study was made.
Reactivity of COS
1.6.3.12 In general it may be stated that a substantial
part of the COS reacts in the Claus reactor.
The extent of reaction of course depends on the temperature
and contact time provided.
Reactivity of CO
1.6.3.13 Although there is negligible reaction
between CO and S02 up to 600°F, it is possible
that CO may react with sulfur to form COS, which in turn could
react with S02 to form sulfur. Since H2S and S02 react very
rapidly, the sulfur vapor formed is available to test the
reaction between CO and sulfur. Therefore a few runs of
1-17
-------
(H-S + S02 + CO) reaction were made. The results of these
runs are reported in Exhibit No. 1-17. These data indicate
that there is negligible reaction between CO and sulfur at
500°F.
Reactivity of H;
1.6.3.14 There is hardly any reaction of H2 with any
other constituent of Glaus feed. Therefore
it may be considered that H2 passes through Glaus process as
an inert.
Intermediate Reactor
1.6.3.15 Carbon monoxide does not react at less than
600°F. Even at 700°F, there is only partial
reaction. There is no reaction of H? in the range of
400-700°F. There is partial reaction of COS in the Glaus
temperature range. These results indicate the necessity of
an Intermediate Reactor to completely react COS, CO and
H2. The actual use of an Intermediate Reactor depends on the
quantity of COS, CO and H2 present in the feed and the
overall economics of the situation.
Effect of COg/HgS Ratio
1.6.3.16 It has been reported in the literature (14)
that the Glaus reaction efficiency is
affected by the C02/H2S ratio in the feed gases. The
(H2S + S02 + H20 + C02) runs were made to test the effect
of C02/H2S ratio on the Glaus reaction efficiency. The
experimental results are reported in Exhibit No. 1-18.
The concentration of C02 was varied from 5% to 2Q% in the
feed gases. The results indicate that C02/H2S ratio has
no effect on the conversion of H2S at 500°F and 1 second
contact time in the range tested.
1.7 THE KINETIC MODEL
1.7.1 General
1.7.1.1 At present the quantitative treatment of
reaction rates rests largely on an empirical
basis, especially for the majority of industrially important
reactions. The interpretation of experimental data and
kinetic analysis is in most cases an individual problem.
1-18
-------
1.7.1.2 The main objective of this program was to
develop a mathematical model for the rate
of the Glaus reactions. It was not the purpose of this
program to go into details of the mechanism of the reaction.
The evaluation of the mechanism of a reaction is a much more
complicated problem than the development of a satisfactory rate
equation. The rate equation may be developed from a knowledge
of the overall reaction and the exact sequence of steps
involved in the reaction need not be known. Therefore
stepwise reactions to produce intermediates were largely
ignored except where the intermediates were detected.
1.7.1.3 For practical reasons it was desirable to
find the simplest model which adequately
simulated the laboratory data over the range of variables.
Chemical kinetics calls for increasing or decreasing the
order of a reactant in the equations, depending upon the
stoichiometry or mechanism. In practice a workable
mathematical model can be developed by using equations that
are first order in each reactant in a majority of the problems
encountered. This is especially true if the number of
equations is large, since the nonlinearity, which may be
necessary to fit the data, is obtained from the large number
of terms. For this reason the model is partially empirical,
but an effort was made to incorporate mechanistic terms so
that when variables are manipulated the computer system will
respond in the same manner as the real system. Use of this
type of approach is justified by the time saved and its
simplicity.
1.7.1.A An effort was also made to abide by theoretical
considerations such as maintaining the proper
ratio of rate constants in a reversible reaction and obtaining
a straight line for the Arrhenius plot.
When a number of reactions take place simultaneously as in
the Glaus process, each reaction may be assumed to take place
at its own specific rate independent of the others and to
follow the simple reaction rate equation. The overall rate
equation may then be considered as the summation of the rates
of all the independent reactions taking place. The general
procedure then consisted of setting up simple differential-rate
equations of the proper order for each separate reaction in
terms of disappearance of reactants and the rate of formation
of products and then to combine them to get the overall rate.
1.7.1.5 The rate constants and the order of reactions
were determined by matching the theoretical
predictions with experimental data on an analog computer.
An assumed rate equation was integrated on an analog computer
to give a relationship between concentration and time. This
theoretical relationship was then compared with the experi-
mental concentration vs. time data to find the k (rate constant)
and the order of reaction.
1-19
-------
1.7.1.6 The search for the values of the rate
constants involved repeated fitting of the
data. Once the best fit to data was obtained at three or
more temperatures, the values of the rate constants were
plotted as log k vs. 1/T. If a straight line or a
reasonable facsimile was not obtained, computer runs were
again made and new values obtained, still maintaining a
reasonable fit to the data with perhaps some of the
deviations from the data now in the opposite direction;
A new Arrhenius plot was drawn and compared with the
previous plot. New straight lines were drawn using all
the points from both trials and the values obtained again
rechecked against the data. By knowing the rate constants
at several temperatures, the activation energy and tempera-
ture dependence of rate were established. The effects of
diffusional resistance were kept to a minimum by using a
high velocity through the catalyst bed.
1.7.1.7 The overall, model was built up stepwise
with an effort made to use the least number
of reactions. The free energy of the reactions and their
equilibrium constants were calculated from values found
in the latest JANAF tables (11). Exhibit No. 1-19 lists
the values of the equilibrium constants at four temperatures.
1.7.1.8 Another consideration was the relative rates
of the reactions. If a reaction rate
constant had a value 10,000 times smaller than the fastest
reaction at any particular temperature, it was ignored,
unless it was involved in a rate controlling step or required
to give correct equilibrium concentrations.
1.7.1.9 The Glaus reaction may be written as:
2 H2S + S02 < > 3/x Sx + 2 H20
The reaction takes place over an alumina catalyst at
temperatures in the range of 392 to 752°F. In the above
reaction Sx is a mixture of the gaseous sulfur species 82,
84, 85 and 83- Below 1650°F, there exists a complicated
equilibrium system between the four species of gaseous
sulfur, which is dependent upon the temperature and the
total sulfur vapor pressure. In the normal temperature Glaus
range So is the most stable and predominent form of sulfur.
It was therefore assumed in this work that gaseous sulfur
exists only as 83-
1.7.1.10 Contact times were calculated based on the
reactor temperature and on the assumption of
507o voids in the catalyst bed. It is given by
_ Void volume of catalyst bed in cu. ft.
Contact Time - fiow rate of feed gases at the reactor tempera-
(sec.) ture and atmospheric pressure in cu. ft./sec.
1-20
-------
1.7.2 Sub-Systems
1.7.2.1 This section deals with the method of
evaluating the rate constants. Kinetic rate
equations for each of the sub-systems were written and analog
programs used to evaluate the rate constants. All the
analog programs were solved on an Electronic Associates Inc.
TR-48 analog computer. An effort was made to abide by the
theoretical consideration of obtaining a straight line for
the Arrhenius plot. Only the forward rate constants could
be determined with any degree of accuracy on the analog
computer. The reverse rate constants were later determined
and incorporated in the overall kinetic model.
H2S + S02
1.7.2.2 The Glaus reaction may be represented by
kl ,
2 H2S + S02 <^ 3/8 S8 + 2 H20
k
In the above reaction k^ is the forward rate constant and
k-[ is the reverse rate constant. The rate constants are
referred to the sulfur dioxide component. It was assumed
that only 83 is formed in the above reaction. Both the
forward and reverse reactions were considered although only
the forward rate constant was determined on the analog
computer. The scaled equations programmed on the analog
computer were:
.5 S02]=0.8 k]_[2.5 S02]fl.25 H2S]-1.6 k^[ 5 Sg][0.3l25 H20]
.25 H2S]=0.8 k]J2.5 S02][1.25 H-SJ-1.6 k^[5 S8H0.3125 H2o]
Ss]=-0.6 kjj2.5 S02][1.25 H2S]+1.2 k-[[5 S8][0.3l25 H20]
^[.3125 H20]=-0.2 kjjl.25 H2S][2.5 S02]+O.A k{[5 Sg][0.3l25 H20]
In the above equations the brackets indicate concentration
of the components. Tha analog schematic is given in
Exhibit No. 1-20.
1.7.2.3 The rate constants for (H2S + S02) reaction
were determined by making several computer
runs and matching the experimental kinetic data with those
predicted by the model. The experimental kinetic data
1-21
-------
reported in Exhibits 1-5 through 1-8 were used for this
purpose. It was found that the rate constants obtained
for (H..S + SO.-..) reaction are approximately the same as
those obtained for (HrS + SO- + HO) reaction. It
therefore follows that the rate of the reaction of
(H2S + S02) is not affected by the 20% water in the
feed.
1.7.2.4 The forward rate constant for the (H2S + S02)
reaction is given in the Arrhenius plot of
log ki vs 1/T°K in Exhibit No. 1-21. The rate constant is
given as a function of temperature by the equation.
1305
k = 61.66 e T
where T is in °K
COS + SQg
1.7.2.5 This reaction may be represented by
k3 .
2 COS + S02 <~> 3/8 Sg + 2 C02,
where k3 is the forward rate constant and k3 is the
reverse rate constant. The rate constants are referred to
the S02 component. It is again assumed that So is mainly
formed in the. above reaction. The scaled equations for the
analog computer are :
COS]= 0.4 k3[8 COS] [5 S0~]-0.1 k:'3[ 20 S8][8 C02]
S02]= 0.125 k3[8 COS][5 SO-]-0.031 k3[20 Sg][8 C021
0 S8]=-0.188 k3[8 COS][5 S02]+ 0.047 k3[ 20 S8][8 C02]
—[8 C02]=-0.4 k3[8 COS] [5 S02]-K).l k3[20 Sg][8 C02]
The analog schematic is shown in Exhibit No. 1-22. The
rate constants were determined by fitting experimental data
with the theoretical predictions. The experimental kinetic
data reported in Exhibit No. 1-11 were used for this purpose.
The Arrhenius plot of log k3 vs 1/T°K is given in Exhibit
No. 1-23. The rate constant may be expressed by the equation
1-22
-------
11420
k3 = 9.5 x 107 e T
where T is in °K
COS + HgQ
1.7.2.6 This reaction is given by
k4
COS + H20 ~^~f H2S + C02
where k4 is the forward rate constant and k4 is the reverse
rate constant. The scaled equations for the analog computer
are:
COS]=2.5 k4[8 COSHO.A H20]- 0.125 k4[8 H2sH8 C02]
4 H20]=0.125 k4[8 COS][0.4 H20]-0.00625 k4[8 H2S][8 C02]
H2S]=-2.5 k4[8 COS][.4 H20]-K).l25 k4[8 H2S][8 C02]
C02]=-2.5 k4[8 COS][0.4 H20]+0.125 k4 [8 H2S][8 C02]
The rate constants were determined by matching the experimental
kinetic data with those predicted by the model. The
experimental kinetic data given in Exhibit No. 1-12 were
used for this purpose. The Arrhenius plot of log k4 vs. 1/T°K
is given in Exhibit No. 1-25. The Arrhenius equation for the
rate may be written as
_ 2880
k4 = 16.87 e T
where T is in °K
COS + SOp + HgQ
1.7.2.7 The reactions involved one given by
^
2 COS + S02 •^~~> 3/8 S8 + 2 C02
k4
COS + H20 <——> H2S + C02
k^
^1
2 H2S + S02 -^—->3/8 S8 + 2 H20
kl
1-23
-------
Using these reactions the analog schematic for (COS + SO.. +
H^O) was developed and is given in Exhibit No. 1-26. The
rate constants for the individual reactions have already
been determined. It was found that the model predicts
the experimental kinetic data reported in Exhibit No. 1-13.
CO + SOg +
1.7.2.8 The reactions taking place may be written
as:
k5
3 CO + S02 < > COS + 2 C02 - (1)
k5
2 COS + S02 < * > 3/8 S8 + 2 C02 - - (2)
ka
COS + H?0 ^ ±~> C02 + H2S - (3)
ki
kl
2 H2S + S02 < > 3/8 Sg + 2 H20 - (4)
kl
In reaction (1) above k5 is referred to the S02 component.
The other rate constants have already been defined. The
scaled equations considering only the forward reactions
are given by
H2S]= 0.32 kL[25 H2S][6.25 S02]-2.5 k4[25 COS][O.A H20]
^{6.25 S02]=0.04 ^[25 H2S][6.25 S02]+ 0.04 k3[25 COS]
[6.25 S02]+0.1 k5[10 CO] [6. 25 S02]
COS ]=2. 5 k4.[25 COS] [.4 H20]+0.32 k3[25 COS] [6. 25 S02]
- 0.4 k5[10 C0][6.25 S02]
20] = 0.04 K4[25 .COSJt.4 H20]- .005 kjj25 H2S][6.25 S02]
C0]= 0.48 k5[10 C0][6.,25 S02]
C02]=-k4[25 COS] [.4 H20]-0.128 k3[25 COS][6.25 S02]
+0.32 k5[10 CO] [6. 25 S02]
1-24
-------
The analog schematic is given in Exhibit No. 1-2?. By
knowing rate constants kj^ k-j and k<^ the rate constant
kc was determined by matching the theoretical predictions
with the experimental data. The experimental data given
in Exhibit No. 1-14 were used for this purpose. The
Arrhenius plot of log kcVs 1/T is given in Exhibit 1-28.
The rate constant equation may then be written
8250
k5 = 3.63 x 105 e T
where T is in °K
1.7.3 The Total System
1.7.3.1 The overall kinetic model was developed
on the basis of the following reactions.
2 H2S + S02 < > 3/8 Sg + 2 H20 - (1)
kl
COS + H20 * > C02 + H2S - (2)
2 COS + S02 3 > 3/8 So + 2 C02 - (3)
<
_ ^
3 CO + S02 s > COS + 2 C02 - (4)
The rate equations may be written as:
7 (H2S) = -2 kiCHzSXSOs) + 2 |^-(S8)(H20) + k4(CO$KH20)
= -k5(CO)(S02) + -(COS)(C02) - k3(COS)(S02)
) - k!(H2S)(S02) + (Sg) (H20)
(COS) = -k4(COS)(H20) + ^-(C02)(H2S) - 2 k3(COS)(S02)
) + k5(CO)(S02) - jJ-(COS)(C02)
1-25
-------
- 3 k5(CO)(S02)+ 3
(H.-.O) = (H; 0) + (H.,S) - (US)
1C 1C
(Sg) = ((S02)IC +(H2S) -I- (COS)C - (S02)-(H2S)-(COS))/8
(C02) = (C02)IC + (CO)IC + (COS^C - (CO) -(COS)
In above equation KSi , KS^, £83, KS5 are the Pseudo-Equilibrium
constants of reaction (1), (2), (3), and (4) respectively.
It may be noted that concentrations of water, sulfur and
carbon dioxide were calculated by material balance.
1.7.3.2 The forward rate constants for the above
reactions have already been determined using
analog programs. The reverse rate constants were determined
in such a manner that the system reaches the theoretical
equilibrium conditions at infinite contact time. The
reverse rate constants were calculated from what may be
called the Pseudo-Equilibrium constant and the method may
be illustrated from the following example
ki
2 H2S + S02 <-^= — > 3/8 Sg + 2 H20
~~
(S02) = -k1(H2S)(S02)+ k{(S8)(H20)
At Equilibrium ^(S02) = 0
.°. k1(H2S)(S02) = k{(S8)(H20)
kl = (S8)(H20)
(H2S)(S02)
Define K5^ = k^/kj^, where KSi is called the Pseudo-Equilibrium
constant. The actual equilibrium constant is related to
Pseudo-Equilibrium constant by
-5/8
_ (H20)eq(S8)eq
K Kb-,
q
1-26
-------
KSi
(H20)eq
By knowing Che actual equilibrium constant and the equilibrium
compositions, the Pseudo-Equilibrium constant may be cal-
culated. Once the Pseudo- Equilibrium constant is known,
the reverse rate constant may be calculated.
1.7.3.3 The overall kinetic model is a digital
program written in Fortran IV. It consists
of subroutines such as the Complex Equilibrium program (12)
and the fourth order Runge-Kutta numerical integration
technique (13). For a given feed composition and temperature,
the program calculates the equilibrium composition of the
various constituents and hence the Pseudo-Equilibrium constant.
From the known values of Pseudo-Equilibrium constant and the
forward rate constants, the reverse rate constants are
calculated. The forward rate constants and the reverse rate
constants of all the four reactions are now known. The
program then integrates the differential rate equations listed
above using the fourth order Runge-Kutta numerical integration.
Then the compositions of the various constituents are
printed out as a function of the contact time. A listing of
the main digital program is given in Exhibit No. 1-30.
This work was done on an IBM 1130 computer.
1.7.3.4 The user must specify the following
information.
(1) System feed composition, mol.%
Feed temperature in °F
Print time and integration interval in seconds.
\-^/
(2)
(3)
The calculated results include
}
)
Reverse rate constants
(2) Equilibrium composition, mol.%
(3) Percent composition of the various
constituents as a function of contact time.
1.7.3.5 The experimental kinetic data of the
overall Glaus reaction have already been
reported in Exhibit No. 1-16. These may now be compared
with those predicted by the model. As a first hand trial
the forward rate constants obtained from the analog
computations were used in the model. In general, the
agreements between the predicted results and the experimental
results were good except for constituents such as COS, CO
and 00%. Therefore the rate constants k3, k^ and k5 were
adjusted to get better agreement between calculated
1-27
-------
and experimental results. The search for these rate constants
was essentially a trial and error process. Several computer
runs had to be made at various temperatures and feed composi-
tions, before arriving at the best values. It was found
that the agreement between calculated and experimental results
is good with the following equations for the rate constants.
k, = 61. 66e
k4 = 2.75e
10870
In the above equations T is in °K. The calculated results
are compared with the experimental values in Exhibit No. 1-29.
The computer printouts are presented as Exhibit No. 1-31.
It may be seen that the agreement between calculated and
experimental results is very good in the normal Claus range.
However at higher temperatures of 700°F, the agreement is
not good. The conversion predicted by the model is less
than those of the experimental values. In fact, the
experimental results show that the equilibrium is violated.
The conversions obtained by the experiment are even better
than what equilibrium permits, which of course is not
possible. More than equilibrium conversions are explained
on the basis that further reaction must be taking place at
lower temperatures (where equilibrium is more favorable)
in the bottom leg and sample tube after the reaction
products leave the catalyst bed at 700°F.
1.7.3.6 From the design point of view the model is
reliable even at 700°F. The model is more
conservative than the experimental values. Equilibrium
limitations have been incorporated in the model. For the
sake of simplicity only the Sg form of sulfur species is
taken into account in the model. This is a reasonable
assumption in the normal Claus range. However other forms
of sulfur such as 85, 8? and 84 also are present at 700°F.
Therefore the model could be improved by taking into
consideration the other forms of sulfur species.
1-28
-------
1.7.3.7 The overall model is quite flexible. It
reduces to the individual reaction models
under individual reaction conditions. Thus we can even
study the kinetics of the individual reactions from the
overall model.
1.7.3.8 In the temperature range of 400 to 700°F,
hydrogen does not react with any constituent
and it passes through the Glaus process as an inert.
Therefore hydrogen was included in the inerts and no
provision was made for printing out of hydrogen composition as
a function of time.,
1.8 BIBLIOGRAPHY
1. Taylor, H. A., and Wesley, W. A., J. Phys. Chem., 31,
216(1926), "The Gaseous Reaction Between Hydrogen Sulfide
and Sulfur Dioxide".
2. Matthews, E., J. Chem. Soc., 2270(1926), "The Interaction
of Sulfur Dioxide and Hydrogen Sulfide",,
3. Lidov, T., "Analiz gazov", p. 74, 1928.
4. Randall, M. and Bichowsky, F., J. Am. Chem. Soc., 40,
368(1918), "Equilibrium in the Reaction Between Water and
Sulfur at High Temperatures. The Dissociation of
Hydrogen Sulfide".
5. Yushkevich, N. F., Karzhavin, V. A., and Avdeeva, A. V.,
Zh. Khim, Prom., 9_, No. 3, 17-26(1932), "The Production
of Sulfur from S02V
6. Lepose, R., Ind. Eng. Chem., 32, 910(1940), "Chemistry
of Sulfur Dioxide Reduction: Kinetics".
7. Gamson, B. W., and Elkins, R. H., Chem. Eng. Prog.,
49_, No. 4, 203(1953), "Sulfur from Hydrogen Sulfide".
8. Munro, A. J. E., and Masdin, E. G., Brit. Chem. Engr.,
1.2, No. 3, 369(1967), "A Study of a Method for
Desulfurizing Fuel Gases".
9. Corrigan, T. E., Chem. Engr., 62_} 200 (April 1955).
10., Weller, S., Am. Inst. Chem. Engrs. J., 2, 61(1956),
"Analysis of Kinetic Data for Heterogeneous Reactions".
11. JANAF Thermochemical Tables, Compiled and edited by the
Dow Chemical Co., Thermal Research Lab., Midland,
Michigan.
1-29
-------
12. Kandiner, H. J., and Brinkley, S. R., Ind. Eng. Chem.,
42, 850(1950), "Calculation of Complex Equilibrium
Relations".
13. Lapidus, L. , "Digital Computations for Chemical Engineers",
McGraw-Hill Co., New York, 1962.
14. Barry, C. B., The Oil and Gas Journal, 63 (May 1970),
"Here's What's Being Done to Combat Sulfur-Oxide Air
Pollution".
1-30
-------
EXHIBIT NO. 1-1
EXPERIMENTAL KINETIC DATA
CATIAYST - POROCEL LPD
SYSTEM - (HgS + SOp + HpO)
FEED - 207. WATER
Contact
Time
Seconds
1
1
Bed
Ht.
Inches
10"
20"
Linear
Velocity
ft/sec.
0.42
0.84
Run
No.
J-16R
J-20R
Nominal
Temp,
°F
500
500
Gas Composition - Volume 70
Feed (Dry Basis)
H2S
2.86
3.0
S02
1.49
1.50
Products
Dry and S-free Basis
H2S
0.39
0.56
S02
0.22
0.25
% Conv. of
H2S
H2/S02-2:1
86
82
7c Conv. of
S02
H2S/S02-2:1
86
83
-------
Stack
UJ
K)
Thermocouples
to TR
Exit
Sample
V A» Heating ~T>
>-P ^
Tape £ |
!. r!
T-
!—
Feed
Sample
Thermocouple;
to TC
i1 nr
i\
i
*
J"
Therm-o-watch
i
*
Water
Saturator
(xj--
Three-Zone
Furnace
Flowmeters
Control Panel
Catalyst Bed
Heating Tape
f I! 11 +1
f f • fit f t t r *- -- - - — -
Bottom Leg
Insulation
EXHIBIT NO. 1-2
Glaus Reactor
Equipment Set-Up
H£S
SO;
CO.
COS CO E£
-------
Exhibit No. 1-3
System - (COS + SQg + HgQ)
Porocel LPD Catalyst
Typical Temperature Profile,
Contact
Time
Sec.
1/2
•*-/ *•
1
2
or
2.5
Bed Ht. from
Bottom of Bed
In Inches
3/4
3 3/4
6 3/4
10 (Top)
3/4
3 3/4
6 3/4
10(Top)
3/4
3 3/4
6 3/4
10 (Top)
Nominal Temperature
400°F (204"C)
210
211
211
212
209
210
210
205
202
206
206
208
475"F (246"C)
245
245
248
250
254
253
255
253
250
248
255
255
550°F (288"C)
288
288
292
285
290
289
289
285
286
288
292
290
700°F (371°C)
370
368
375
370
370
376
383
370
368
368
374
370
r
LO
LO
-------
EXHIBIT NO. 1-4
System - HgS + SOg
Porocel LPD Catalyst
Typical Bed Temperature Profiles.,
Contact
Time
Sec.
1/16
1/8
1/8
1/4
Bed Ht. from
Bottom of Bed
in Inches
1/4
3 (Top )
1/4
3 1/4
6 (Top)
1/4
3 (Top)
1/4
3 1/4
6 1/4
10 (Top)
Nominal Temperature
450°F(232°C)
236
234
230
232
234
228
500°F(260UC)
263
259
260
264
260
255
550UF(288UC)
287
285
308
300
280
291
303
317
282
700aF(371wC)
368
373
378
375
363
370
375
395
387
-------
EXHIBIT NO. 1-5
EXPERIMENTAL KINETIC DATA
CATALYST - PORQCEL LPD
SYSTEM - HgS + S02
Contact Time
sec.
1/2
1
2
1/2
1
2
1/2
1
2
Run No.
7
8
9
2
1
3
6
4
5
Nominal
Temp.
°F
400
400
400
550
550
550
700
700
700
Gas Composition - Volume 70
Feed
H2S
11.65
12.5
11.9
1.4
1.94
1.89
5.93
6.38
6.22
S02
1.8
2.05
1.88
10.9
10.15
11.00
3.15
3.05
3.09
Products (Dry and S-Free Basis)
H2S
9.6
10.83
8.93
trace
trace
trace
0.72
0.73
0.69
S02
1.13
1.22
0.56
10.4
9.26
9.93
0.64
0.5
0.67
I
u>
Ln
-------
EXHIBIT NO. 1-6
Experimental Kinetic Data
Catalyst - Porocel LPD
System - H2S + S02
Contact
Time
Sec.
1/8
1/4
1/8
1/4
1/2
1/16
1/8
1/4
1/16
1/9
1/4
Run
No.
F-4
*E-1
: F-3
"D-2
D-l
F-l
2AA
2B
F-2
2A
6B
Bed
Ht. in
Inches
3
10
3
10
10
3
6
10
3
6
10
Nominal
Temp.
°F
450
450
500
500
500
550
550
550
700
700
700
Gas Composition - Volume 70
Feed
H2S | S02
6.04
5.77
5.33
6.05
5.85
5,76
5.6
6.36
4.84
5.99
6.59
3.06
2.86
2.74
2.92
3.0
2.77
2.65
3.23
2.6
2.9
3.3
Products (Dry and S-free Basis)
H2S
3.16
3.64
0.78
2.07
.67
1.86
0.84
0.34
1.17
0.91
0.89
SO*
1.68
1.95
0.63
1.13
0.68
0.84
0.36
0.33
0.85
0.34
0.44
u>
* It was found after Runs E-l and D-2 that liquid sulfur
had deposited on the catalyst which probably could
have poisoned the catalyst.
-------
EXHIBIT NO. Ir7
Experimental Kinetic Data
Catalyst - Porocel LPD
System - (H?S + SOg)
Contact
Time,
sec.
1/2
1
1/2
1
Run
No.
P-l
P-2
P-3
P-4
Nominal
Temp. ,
OF
400
400
450
450
Gas Composition - Volume %
Feed
HpS
2.94
2.69
2.97
2.7
SO?
1.55
1.40
1.47
1.39
Products
(Dry and S-Free Basis)
H^S
0.33
0.0
0.075
0.06
SO^
0.26
0.11
0.0125
0.00
1-37
-------
EXHIBIT NO. 1-8
I
u>
00
Experimental Kinetic Data
Catalyst - Porocel LPD
System - (HgS + SOg + HP0)
Contact
Time
Sec.
1
2.5
1/2
2.5
1/2
2.0
1/16
1/16
Run No.
J-4
J-5
J-7
J-6
J-l
J-8
*J-9
*J-10
Nominal
Temp.
°F
400
400
450
450
500
500
550
700
Gas Com
Feed (Dry Basis)
H2S
2.45
2.2
2.62
2.08
6.77
2.74
2.88
3.19
S02
1.17
1.0
1.08
0.83
3.73
1.13
1.55
1.61
position - Volume %
Products (Dry and S-free Basis)
H2S
0.16
0.45
0.51
0.7
0,52
1.05
0.96
1.28
S02
0.06
0.025
0.03
0.04
0.60
0.23
0.66
0.56
* Runs J-9 and J-10 are made with 3" of bed.
-------
Exhibit No. 1-9
Experimental Kinetic Data
Catalyst - Porocel LPD~~
System - (Hg + SOg + HgO)
Feed - 20% Water
Contact
Time,
Sec.
2.5
2.0
1/2
1
2.0
Run
No.
M-l
M-2
M-3
M-4
M-5
Nominal
Temp.
°F
400
550
700
700
700
Gas Composition - Volume 70
Feed (Dry Basis)
H2
1.31
1.19
0.83
1.27
1.30
SOp
1.27
1.23
1.38
1.44
1.08
Products (Dry and S-Free Basis)
E2
1.28
1.18
0.83
1.02
1.25
S02
1.22
1.18
1.38
1.34
1.08
H£S
0.0
0.0
0.0
0.0
0.09
I
u>
SO
-------
EXHIBIT NO. 1-10
Experimental Kinetic Data
Catalyst - Porocel LPD
System - (COS + Na)
Contact
Time
Sec.
1
2
0.554
2.22
1/2
1
Run No.
H-4
H-3
22R
H-2
H-5
H-6
Nominal
Temp.
°F
400
400
550
550
700
700
Gas Composition - Volume 70
Feed
COS
2.04
1.82
2.26
2.05
1.92
1.81
Products (Dry and S-free Basis)
COS
1.81
1.39
1.78
1.33
1.22
1.11
C02
0.1
0.14
0.38
0.25
0.31
0.31
CS?
0.24
0.38
0.26
0.59
0.50
0.45
CO
Trace
0.01
0.01
0.01
0.02
0.03
I
.p>
o
-------
EXHIBIT NO. 1-11
Experimental Kinetic Data
Catalyst - Porocel LPD
System - (COS + SO?)
Contact
Time
Sec.
1
2.0
1
2
1/2
1
Run
No.
S-5
S-6
S-lR
S-2R
S-3
S-4
Nominal
Temp.
°F
400
400
550
550
700
700
Gas Composition - Volume %
Feed (Dry-Basis)
COS
0.79
0.61
0.71
0.47
0.87
0.69
S02
1.39
1.38
1.49
1.41
1.61
1.47
Products
COS
0.78
0.60
0.45
0.21
0.02
0.0
Dry and S-Free Basis)
SOp
1.39
1.39
1.35
1.29
1.14
1.1
COp
< .01
< .01
0.18
0.21
0.62
0.51
-------
Exhibit No. 1-12
Experimental Kinetic Data
Catalyst - Porocel LPD
System - (COS + "
Feed - 20% Water
Contact
Time
Sec.
1/2
1
2.5
1/2
1
2
1/2
1
2
Run No.
N6
N7
Nl
N8
N9
N2
N3
N4
N5
Nominal
Temp.
°F
400
400
400
550
550
550
700
700
700
Gas Composition - Volume %
Feed (Dry-Basis')
COS
0.9
0.94
0.94
0.87
0.86
0.88
0.89
0.89
1.04
Products
COS
0.37
0.12
0.02
0.27
0.12
0.01
0.04
0.04
0.0
Dry and S-Free Basis)
C02
0.49
0.61
0.66
0.50
0.56
0.78
0.72
0.61
0.8
RpS
0.46
0.61
0.59
0.53
0.62
0.89
0.83
0.67
0.9
I
4S
NS
-------
Exhibit No. 1-13
Experimental Kinetic Data
Catalyst - Porocel LPD
System - (COS + SOg + HgO)
Feed - 20% Water
u:
Contact
Time
Sec.
1/2
1
2-5
1/2
1
2.5
1/2
1
2
1/2
1
2
Run
No.
K- 7
K- 8
K- 2
K-13
K-12
K-ll
K- 9
K-10
K- 3
K- 4
K- 5
K- 6
Nominal
Temp.
°F
400
400
400
475
475
475
550
550
550
700
700
700
Gas Composition - Volume %
Feed (Dr~s
COS
0.9
0.945
0.79
0.89
0.88
1.09
0.83
0.85
0.97
0.89
0.91
0.98
r-Basis)
S02
1.49
1.67
1.19
1.50
1.55
1.18
1.48
1.38
1.20
1.47
1.45
1.12
Products (Dry and S-Free Basis)
COS
0.59
0.33
0.01
0.50
0.2
0.06
0.35
0.1
0.0
0.035
0.02
0.0
S02
1.29
1.21
0.2
1.38
1.45
0.79
1.30
1.1
0.9
1.25
1.37
0.94
C02
0.19
0.43
0.79
0.29
0.54
0.78
0.47
0.61
0.8
0.76
0.79
0.85
H2S
0.02
0.035
0.13
0.11
0.22
0.17
0.04
0.18
0.28
0.64
0.64
0.68
-------
Exhibit No.1-14
Experimental Kinetic Data
Catalyst - Porocel LPD
System - (CO + SOg + H30)
Feed - 20% Water
Contact
Time,
Sec.
1
2.5
1.0
2.0
1.0
2.0
1/2
1
2
Run No.
L-6
L-l
L-7
L-2
L-9
L-8
L-3
L-4
L-5
Nominal
Temp.
°F
400
400
550
550
600
600
700
700
700
Gas Composition - Volume %
Feed (Dry Basis')
CO
0.79
0.93
0.72
0.83
0.69
0.71
0.8
0.76
0.85
S02
1.56
1.27
1.51
1.30
1.51
1.43
1.4
1.38
1.08
Products (Dry and S-Free Basis)
CO
0.73
0.76
0.71
0.76
0.46
0.39
0.68
0.56
0.46
S02
1.52
1.44
1.51
1.28
1.44
1.25
1.38
1.3
0.9
C02
0.0
0.0
0.0
0.06
0.13
0.27
0.08
0.15
0.34
COS
0.0
0.0
0.0
0.01
0.01
0.016
0.01
0.01
0.0
H?S
0.0
0.0
0.0
0.0
0.01
0.04
0.01
0.02
0.06
I
.p-
.p-
-------
Exhibit No. 1-15.
Experimental Kinetic Data
Catalyst - Porocel LPD
System - CO + HgO
Feed - 20% Water
Contact Time
Sec.
2.0
1.0
2.0
Run No.
Q-2
Q-3
Q-4
Nominal
Temp.
°F
550
700
700
Gas Composition - Volume %
Feed ("Dry-Basis)
CO
0.80
0.66
0.76
Products (Dry and S-Free Basis)
CO
0.79
0.68
0.75
Ui
-------
EXHIBIT NO. 1-16
EXPERIMENTAL KINETIC DATA
CATALYST - POROCEL LPD
SYSTEM - OVERALL REACTION
FEED - 207. WATER
Contact
Time
Sec.
1/4
1/2
1
2
1/8
1/2
1
2
1/8
1/2
1
2
Run
No.
V-12
V-7
V-8
V-9
V-ll
V-l
V-2
V-3
V-10
V-4
V-5
V-6
Nominal
Temp.
°F.
400
400
400
400
550
550
550
550
700
700
700
700
Gas Composition - Volume %
Feed (Dry - Basis)
CO
0.98
0.98
0.84
0.83
0.88
0.95
0.78
0.86
0.92
0.98
0.78
0.85
COS
0.98
0.97
1.01
0.82
0.95
0.99
0.95
0.85
1.03
1.07
0.90
0.82
H2
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
H2S
1.56
1.38
1.57
1.61
3.33
3.67
3.19
3.48
3.92
3.17
3.17
3.38
S02
3.10
3.31
2.42
2.42
3.60
3.60
2.99
3.16
3.91
3.23
3.01
3.48
C02
6.7
5.23
6.85
6.5
6.2
6.46
7.07
5.7
6.39
6.12
6.92
5.63
CO
0.98
0.98
0.84
0.83
0.88
0.95
0,78
0.81
0.85
0.85
0.68
0.54
Products (Dry and S-free Basis)
COS
0.72
0.63
0.62
0.13
0.71
0.39
0.23
0.0
0.36
0.08
0.04
0.0
H2
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
0.8
H2S
0.26
0.04
0.0
0.03
1.35
0.31
0.40
0.3
1.59
1.05
1.28
1.47
S02
2.06
2.55
1.33
1.08
2.9
1.42
1.19
0.97
2.16
1.96
1.48
1.33
C02
6.96
5.37
7.17
6.84
6.4
6.99
7.66
6.48
6.77
6.89
7.71
6.91
-------
EXHIBIT NO. 1-17
Experimental Kinetic Data
Catalyst - Porocel LPD
System - H?S + SOp + CO
Contact
Time
Sec.
1
2
Run
No.
c-i
C-2
Nominal
Temp.
°F
500°F
500°F
Gas Composition - Volume 70
H2S
3.14
3.25
S02
5.41
4.88
Feed
CO
6.45
5.53
COp
.05
.04
COS
.01
.01
CS2
ND
ND
Products (Dry 4 S-Free Basis)
H2S
ND
ND
SOp
4.66
3.41
CO
7.34
5.84
COp
.07
.07
COS
.02
.02
CS2
ND
ND
ND - None Detected
-------
EXHIBIT NO. 1-18
Experimental Kinetic Data
Catalyst - Porocel LPD
System - (HPS +SQg + HgO + COg)
_
Feed - 20% Water
Contact
Time
Sec.
1
1
1
1
Run
No.
J-16
J-17
J-18
J-19
Nominal
Temp.
°F
500
500
500
500
Gas Composition - Volume %
Feed (Dry-Basis)
H2S
2.76
2.97
2.89
2.92
S02
1.38
1.41
1.37
1.39
C02
0
5.14
10.39
19.06
Products (Dry and S-free Basis)
H2S
0.34
0.68
0.38
0.42
S02
0.13
0.19
0.15
0.13
C02
0
5.14
10.39
19.06
Conv.
of
H2S for
Ikl = 2:1
S02
88
81
92
90
Conv.
of
S02 for
H&S - 2-1
so2 • L
91
87
89
91
-------
EXHIBIT .L~19
, Eqm. Const.
2 H2S + SC^ < > 3/8 Sg + 2 H20
3 CO + S02 > COS + 2 C02
2 CCS + SC^ — > 3/8 Sg + 2 C02
CCS + 1^0 ^ > HgS + CC^
f*f\ J_ tf ^ •^ IT ^L f*f\
\^\J i^ ^*O ^ ^^^^B^.^™— ^*O ^^ wWO
2 COS ^Z^ 2 CO + 2/8 Sg
2 COS ^1^ CS2 + C02
2 H2 + S02 > 1/8 Sg + 2 H20 -
2 C$2 + Sfy > 3/8 Sg + 2 COS
CS2 + 2 H20 > 2 H2S + CC^
261°F
400°K
6.9 xlO7
1.85xl028
3.44xl016
2.23xl04
1.55xl03
0.33x 10'8
0.225
4.4xl018
6.78x 1017
2.21X 109
440 °F
500°K J
8.87 xlO5
3.1xl021
9.86X1011
2.86xl03
138
0.47xlO"6
0.229
1.12X 1014
1.89x 1013
3.57x 107
682°F
600 °K
1.86X103
2.05xl016
7.83x10°
719
28.4
O.UxlO'4
0.231
9.14xl010
1.81x 1010
2.24xl06
800°F
700°K
95.5
4.14xl012
6.97xl06
270
9.55
0.14xlO"3
0.231
5.4xl08
1.31x 108
3.16x 105
-------
EXHIBIT NO. 1-20
Analog Schematic
(HgS + SOa) Reaction
— ° -/ST
TP^
46:
1-50
-------
30
20
2 H2S
T
T
EXHIBIT NO. 1-21
Arrhenius Plot
S0
kl
,
kl
3/8 S8 + 2
10
8
o
01
w
I
o
>
1.4
700
1.5
1.6
600
I
1.7
T°K
x 103
5flOi
500
1.8
1.9
450
2.0
40i
°F
2.1
1-51
-------
EXHIBIT NO. 1-22
+ -0/*,[8Cos
2/
I
8
-K/4
Analofi Schematic
(COS + SO?) Reaction
+/>cc*7
/>o s,7
•»•* -*•/
-y
MULT
6
U
-y
Ml/LT
1-52
-------
10
7
T
T
EXHIBIT NO. 1-23
Arrhenius Plot
3 .
2COS + S02 < - ~ 3/8 S8 + 2 C02
1
.7
.4
.2
.1
.07
.04
.02
.01
007
o
-------
EXHIBIT NO- 1-24
Analog Schematic
(COS + HrO) ReacEion
-/o
1-54
-------
10
7
EXHIBIT NO. 1-25
Arrhenius Plot
T T
COS +
H2S + C02
i
.7
.4
.2
.1
.07
.04
.02
.01
1.4
o
-------
EXHIBIT NO. 1-26
Analog Schematic
(COS + SQg + HP0) Reaction
1-56
-------
-10
Hi Atl ID .Li. IXV- *.-«-*
Analog Schematic
(CO + SOg + HaO) Reaction
1-57
-------
1
.7
.4
.2
.1
.07
.04
.02
.01
.007
.004
.002
o
0)
CO
o
EXHIBIT NO. 1-28
Arrhenius Plot
3 CO + SDo^ > COS + 2 CO..
.001
0007
0004
0002
0001
1.4
700
L_
600
550
500
450
400*
1.5
1.6
1.7
x 103
1.8
1.9
2.0
2.1
1-58
-------
COMPARISON OF EXPERIMENTAL RESULTS WITH
THOSE PREDICTED BY THE MODEL
400°F, 1/2 sec.
Feed
Expt.
Model
CO
0.98
0.98
0.98
Gas Composition - Volume %
COS
0.97
0.63
0.67
H2
0.8
0.8
0.8
H2S
1.38
0.04
0.06
S02
3.31
2.55
2.51
C02
5.23
5.37
5.57
400°F. 1 sec.
Feed
Expt.
Model
0.84
0.84
0.84
1.01
0.62
0.47
0.8
0.8
0.8
1.57
0.0
0.07
2.42
1.33
1.4
6.85
7.27
7.44
400°F, 2
Feed
Expt .
Model
0.83
0.83
0.83
0.82
0.13
0.18
sec.
0.8
0.8
0.8
1.61
0.03
0.07
2.42
1.08
1.33
6.5
6.84
7.2
550°FJ 1/2 sec.
Feed
Expt.
Model
0.95
0.95
0.94
0.99
0.39
0.40
0.8
0.8
0.8
3.67
0»31
0.55
3.6
1.42
1.75
6.46
6.99
7.16
550°F, 1 sec.
Feed
Expt.
Model
0.78
0.78
0.77
0.95
0.23
0.17
0.8
0.8
0»8
3.19
0.40
0.62
2.99
1.19
1.32
7.07
7.66
7.96
550°F, 2 sec.
Feed
Expt.
Model
0.86
0.81
0.82
0.85
0.0
0.03
0.8
0.8
0.8
3.48
0.3
0.66
3.16
0.97
1.33
5.7
6.48
6.65
700°F, 1/2 sec.
Feed
Expt.
Model
0.98
0.85
0.89
1.07
0.08
0.05
0.8
0.8
0.8
3.17
1.05
2.08
3.23
1.96
2.13
6.12
6.89
7.29
700^, 1 sec.
Feed
Expt.
Model
0.78
0.68
0.65
0.9
0.04
0.01
0.8
0.8
0.8
3.17
1.28
2.16
3.01
1.48
2.0
6.92
7.71
8.0
700°F, 2 sec.
I feed
Expt.
1 Model, ,
0.85
0.54
O.S6
0.82
0.0
0.1
0.8
0.8
0.8
3.38
1.47
2.1
3.48
1.33
2.3
5.63
6.91
6.79
1-59
-------
111)
// JOU
/ / F OR
•lOCSCCARDi TYPLWRI T L'R * 1 i 32PR I NTER .KEYBOARD, U I S< )
"EXTENDED PRECISION
•ONE WORD INTEGERS
*L1ST ALL
DEFINE FILE 11(50 .100.U.I 11)
REAL KK,KF1 ,KF2,KF3»KR,KR1.KR2»KR3.K(6),KP.M(11)
DIMENSION A(6.7)
COMMON A
CALL INP
GO TO 20
1 ICT=1
CALL DATSWll»IPR1)
GO TO (20.2D.IPR1 EXHIBIT NO. 1-30
200 FORMAT!1I TLMP )')
20 WRITE!1.200)
READ!6, 10)T
TK = (T-3?.J/l.fl+273.1
KF = 61,.66*EXP(-1305./TK)*1.
KF1 = 2.75«EXP(-2080./TK)*l.t
KK2 = 2.09E+07*KXP(-10870./TK)
KF3 = rtl3.#CXP(-6580../TK)*l.t+0'+
ICT = 2
21 CALL DATSW12,IPR2)
GO TO (22,23") ,IPR2
201 FORMAT!M S02 >( H2S )( SH M H2»
1 )( COS )')
C-UiJ
22
601
23
10
202
24
50
51
300
WRITE! 1,201)
READ! 6, 10)S02.H2S.S8.H20,C02,CO.S
FORMAT!'( CO )( H2 )')
WRITE(1,601)
READ(6.10)CO,H2
ERT = 100.-SO? -H2S-S8-H20-C02-COS-CO-H2
ICT = 2
CALL DATSWJ3.IPR3)
GO T0124.50)*1PR3
FORMAT!1X,6(F10.2,2X))
FORMAT!'( TIME
WRITE!1,202)
READ!6.10)TIMK,AII
ICT = 2
GO TO (25,51),ICT
DO 300 :i = l,ll
) ',' ( INTO INT ) ' )
M ( I )=0..0
K 1) =S02
M 2) =H'2S
•-1 5) =S8
V 6) =H20
V 7) =C02
M 8) =COS
K 9)=CO
M 10)=H2
M 11)=ERT
READ!11'1)N,NC,NJJ
WRITE!11'N+3)M
CALL VAR8
DO 361 1=1,11
361 MF(IJ=M II)
1-60
-------
CALL EU (T(MF.K)
MF(5) = MF(5l+MF(4l«6./8.+MF<31*2.
PAuE 02
"F(3)=0.
TOT =0.0
DO 362 1=1.11
362 Es;: ;?i;T::i:w,s,»-,,./.../-F,i./MF,2.«2,.. TOT,*,*.,.., EXHIBIT »• ^ (^1—>
HP = EOH«MF(2)»MF(5)»»l5./8.)/IMF(6)«TOT«(5./tt. ) )
KR"KF/ICP
EOK1 = MF(2)*MF(7)/(MF( 8)»MF<6>)
KR1 • KF1/EOK1
EOK2 » (MF(7)««2»MFl5)«»13./8.)I/IMF!1)«MFI 8)*»2)»TOT««(5./6.I
HP = EQK2»MF( 8>**F(5)»*<5./8. I/ I MF I 7 ) *TOT*« ( b. / 8. ) )
KR2 = KF2/KP " ' .
EQK3 = (MF18)«MF(7)»«2«TOT)/IMF(9)**3*MF(1) )
KP=EOK3»MF(9)/MF(7)
DO 364 1 = 1.11
364 MF( I )a!4F< I )/TOT«100.
WRITEI3.101)
WRITE13.106)
WRITEI3.321)
WRITE!3.103) MF<1).MFI2IiMF(S).MF16).MF(7),MF(8I.MKI9).MKI 10)iriF(
111)
321 FORMAT!IX,•»»»TIME»»«'7X.'S02'»7X.'H2S1.6X,'S8'.7X.'M20'.7X«'CO21•
17X.-COS'.8X.iCO'.8X.'H2'.5X.' INERT1/)
RES=0.0
WRITE!3.102 IRES.S02iri2S.S8iH20.C02.COSiCO.M2.tKT
25 TIM1=TIME
CALL RKG5 (M ,MF.T,TIMl.RES.KF.KFl.KH2.KF3.Kr(.fCKl,i(K2.KKJ.Al I )
S=0.0
DO 60 1 = 1 ill
M (D-MFIII
60 S=S+viF( I I
DO 61 1=1.11
61 VFII)=MFII)/S«100.
RES =RES * .00001
WRITE I 3. 100) RES..IF I 1 ) ,*F I 2 ) , MF [ 5 J . -in 6 ) ,,/,F( 7) ..-.Fld) ,,-,K(9) ,.1H I 11)
RES = HES-.00001
GO TO 1
100 FORMAT! lX.8FlJ.4ilJAiK10.4l
101 FORMAT!iMl)
102 FORMATl1X.10F10.4)
106 FORMAT! IX, ' TE^. '• 11 J . 5 . ' JEG. r ' . / 1 X . • t JO IL .
11X.'FORftAKD K ' .4E13.5/1A. ' BAC->2 =ou/j n25 =OJOJ So =Jo3j ~>e.-J =JOob
ECOS "nO^C CO =JOHF n2 =D,T92 E: I ' =JjCV
.%J =OOC^ .\C =OOC9 SJJ =o:.'CA
STATEMENT ALLOCATIONS
230 =0115 201 =ail-J 6>01 =0143 1J =Ulbl ^J2 sJl1?/ i^i =-i6& ijj =. i v i iui =j±tf »-i 'JiirA l^o =Cl-»t
1-61
-------
TEMP. 0.40000E
EQUIL. K 0.39510E
FORWARD K 0.40000E
BACKWARD K o.
...TIME".
EQUIL.
0.0000
0.1000
0.2000
0.3000
0.4000
0.5000
0.6000
0.7000
0.8000
0.9000
TOL. ON EQUIL.
0.17866E-24
0.17876E-24
51023E
S02
1.3218
3.3100
2.6861
2.5996
2.5629
2.5336
2.5071
2.4825
2.4598
2.4386
2.4191
03 PEG. F
06 0.42765E 04 0.72260E 13
05 0.35000E 03 0.25:>0dt 02
03 0
0.
1.
0.
.81841E-01 0.74376£-04
H2S
1163
3800
1820
0.0731
0.
0.
0.
0.
0.
0.
0.
0606
0589
0585
0583
0580
0578
0576
S8
0.5383
0.0000
0.2387
0.2718
0.2858
0.2969
0.3070
0.3163
0.3250
0.3330
0.34Q4
H23
22.3962
20.0000
21.2830
21.4038
21.4213
21.4269
21.4310
21.4346
21.4380
21.4411
21.4440
0.62252E 23
0.79999t :;i
O.V3729E-13
7
5
5
5
5
5
5
5
5
5
5
C0<
.2876
.2300
.3214
.3913
.4548
.5135
.5680
.6185
.6653
.7087
.7490
COS
O.OUUu
0.9700
0.9039
0.03BO
0.7766
0.7196
0.6669
0.6179
0.5726
0.53u6
0.4916
0
0
0
0
0
0
0
0
0
0
0
CJ
.UWv,
•9tiOu u.t
.9832
.9831
.9827
.9323
.9819
.9814
.9610
«9dO:>
.93J1
.-
; — ,
CONSTANT NOT MET
0.
0.
25532E
25525E
15
15
0.21649E
0.22357E
03 0.
03 0.
65435E 12
69766E 12
0.27234E
0.27234E
13
13
0- ??r)99E
0.22100E
21
21
EXHIBIT NO. 1-31
bo. 33:
67.
67,
67,
6/,
67.66Jd
6 7 . o 7 3 2
67.bead
o/,
67.70'
•07.71:
1-62
-------
TEMP.
EQUIL. K
FORWARD K
BACKWARD 1C
•«»TIME»»«
EOUIL.
0.0000
0.1000
0.2000
0.3000
0.4000
0.5000
0.6000
0.7000
o.eooo
0.9000
1.0000
1.1000
1.2000
1.3000
1.4000
1.5000
1.6000
1.7000
1.8000
1.9000
2.0000
2*1000
0.40000E 03 DEC. F
0.39311E 06 0.42764E 04 0.72256E 13 0
0.40000E 05 0.35000E 03 0.25000E 02 0
0.31527E 03 0.81844E-01 0.46078E-04 0
S02
0.4913
2.4200
1.7756
1.6380
1.5858
1.5546
1.5300
1.5084
1.4887
1.4705
1.4536
1.4380
1.4235
1.4101
1.3977
1.3861
1.3754
1.3655
1.3563
1.3478
1.3399
1.3325
1.3257
M2S
0.1914
1.6100
0.3672
0.1458
0.0929
0.0788
0.0746
0.0730
0.0720
0.0713
0.0707
0.0701
0.0696
0.0691
0.0686
0.0682
0.0677
0.0673
0.0670
0.0666
0.0663
0.0660
0.0657
S8
0.5295
0.0000
0.2452
0.2976
0.3174
0.3293
0.3386
0.3468
0.3542
0.3611
0.3675
0.3734
0.3789
0.3840
0.3887
0.3931
0.3971
0.4008
0.4043
0.4075
0.4105
0.4133
0.4158
H2O
22.5398
20.0000
21.3310
21.5713
21.6314
21.6498
21.6574
21.6620
21.6636
21.6689
21.6718
21.6745
21.6770
21.6794
21.6816
21.6836
21.6855
21.6873
21.6889
21.6905
21.6919
21.6932
21.6944
.98742E 23
•79999E 01
.50642E-13
C02
8.2668
6.5000
6.5857
6.6479
6.7028
6.7530
6.7994
6.6423
6.8U20
6.9189
6.9530
6.9647
7.0140
7.0412
7.0664
7.0897
7.1114
7.1314
7.1500
7.1672
7.1832
7.1980
7.2118
COS
0.0000
0.8200
0.7645
0.7090
0.6569
0.6085
0.5636
0.5220
0.4634
0.4477
0.4147
0.3840
0.3557
0.3294
0.3051
0.2826
0.2617
0.2424
0.2245
0.2079
0.1926
0.1784
0.1652
Co
0.0000
0*8300
0.8329
0.8333
0.8333
0.8331
0.8329
0.8328
0.8326
0.8324
0.8322
0.8319
0.8317
O.H315
0.8313
0.8311
0.8309
0.8306
0.8304
0.8302
0.8300
0.8297
0.8295
H2 iNtuT
O.OJOU 67.9609
0.8000 67.0200
67.2940
67. 3526
67.3748
67.30U1
67.3VH6
67.4078
67.41o2
67.<.^«0
67.4312
67.4378
67.4440
67.449B
67.4551
67.4600
67.4646
67.46BU
67.4728
67.4764
67.4798
67.4B29
67.4859
TOL. ON EOUIL. CONSTANT NOT MET
0.17866E-24 0.25532E 15 0.21649E 03 0.65435E 12 0.27234E 13 0.22099t 21
0.17914E-24 0.25498E 15 0.21074E 03 0.61924E 12 0.27234k. 13 0.22099E 21
1-63
-------
TEMP. 0.40000E 03 DEG. F
EOUIL. 1C 0.39509E 06 0.42764E 04 0.72256E 13 0
FORWARD 1C 0.40000E 05 0.350COE C3 3.250COE 02 0
BACKWARD <
...TIME...
EOUIL.
0.0000
0.1000
0.2000
0.3000
0.4000
0*5000
0.6000
0.7000
0.8000
0.9000
1.0000
1.1000
1.2000
1.3000
1.4000
1.5000
1.6000
1.7000
1.8000
1.9000
2*0000
2*1000
2*2000
2.3000
2.4000
2.5000
2.6000
2.7000
2.8000
2.9000
3.0000
3.1000
3.2000
0.42752E 03 0.81843E-01 0.62482E-04 0
S02
1.0220
3.1000
2.4124
2.3070
2.2655
2*2349
2.2076
2.1826
2.1594
2.1379
2.1180
2.0995
2.0823
2.0665
2.0517
2.0381
2.0254
2.0137
2.0028
1.9927
1.9834
1.9747
1.9667
1.9592
1.9522
1.9458
1.9398
1.9343
1.9292
1.9244
1.9199
1.9158
1.9120
1.9084
H2S
0.1349
1.5600
0.2360
0.0901
0.0691
0.0657
0.0648
0.0643
0.0639
0.0635
0.0632
0.0629
0.0626
0.0623
0.0620
0.0617
0.0615
0.0613
0.0611
0.0609
0.0607
0.0605
0.0604
0.0602
0*0601
0*0600
0*0599
0.0598
0.0597
0.0596
0.0595
0.0594
0.0593
0.0593
S3
0.5713
0.0000
0.2628
0.3031
0.3188
0.3305
0.3408
0.3503
0.3592
0.3673
0.3749
0.3819
0.3884
0.3944
0.4000
0.4051
0.4099
0*4143
0.4184
0.4222
0.4257
0.4290
0.4320
0*4348
0*4374
0.4398
0*4420
0*4441
0*4460
0.4478
0.4494
0.4510
0.452't
0.4537
H20
22.5727
20.000U
21.4184
21.5788
21.6055
21.6132
21.6178
21.6218
21.6254
21.6287
21.6318
21.6347
21.6373
21.6398
21.6421
21.6443
21.6463
21.6481
21.6498
21.6514
21.6529
21.6543
21.6555
21.6567
21.6578
21.6588
21.6598
21.6607
21.6615
21.6622
21.6629
21.6636
21.6642
21.6648
.62573E 23
.79999E 01
.81898E-13
C02
8.7946
6.7000
6.8008
6.8734
6.9384
6.9984
7.0539
7.1054
7.1531
7.1973
7.2382
7.2762
7.3114
7.3440
7.3743
7.4023
7.4283
7.4524
7.4747
7.4955
7.5147
7.5325
7.5490
7.5644
7.5786
7.5918
7.6041
7.6155
7.6261
7.6359
7.6450
7.6535
7.661'i
7.6687
CJi
O.OOJ-J
0.9800
0.9134
0.8465
0.7841
0.7202
0.6725
0.6228
0.576 /
0.5341
0.49*6
0.45o 1
0.4242
0.3929
Or 3*38
0.3369
0.3121
0.2890
0.2677
0.2479
0.2296
0.2127
0.1970
0.1825
0.1691
0.1567
0.1451
0.1345
0.1246
0.1155
0.1070
0.0992
0.0920
0.0653
Co
O.JJGOJ
0.9800
0.9336
0.9837
0.9934
0.9831
0.9827
0.9824
0.9820
0.9816
0.9di 3
j.9509
0.9dO-
0.9801
0.9797
0.9793
0.9789
0.97d5
0.9781
0.9777
0.9773
0.9768
0.9764
0.9760
0.9756
0.9752
0.974d
0.9743
0.9739
0.97 35
0.97.51
0.9726
0.9722
0.9710
0.8COU
66. 9u».2
65.3600
06. lo«7
66.2130
66.2303
66.2O3*
60 « £ 7s£
oo • 2 s*« 3
66 . 3J^3
60.3077
66. 3144
66.3^64
66.3318
66.3367
66.3413
66.3456
66.3496
60.3333
66.3567
66.3599
66.3626
66.36 30
60.3ob 1
66.3705
66.3727
66.3747
66.3?o4
66.3600
66 . joi6
1-64
-------
TEMP, 0.55000E 03 DEG. F
EOU1L. K 0.80092E 0% 0.1172SE 04 0.11017E 11 0.13061t 19
FORWARD K 0.60000E OS 0.67000E 03 0.780COE 03 0.65000t 02
BACKWARD K 0.33256E 04 0.57126E 00 0.36862E-01 0.33899E-09
EOUIL.
0.0000
0.1000
0.2000
0.3000
0.4000
0.5000
502
0.9537
3.6000
1.9834
1.8753
1*8231
1.7819
1.7*79
H2S
1.1018
3.6700
0.5764
0.5029
0.5175
0.5342
0.5487
S8
0.7953
0.0000
0.6195
0.6606
0.6802
0.6956
0.7082
H20
23.8378
20.0000
23.3383
23.4282
23.4216
23.4113
23.4020
C02
8.5609
6.4600
6*7027
6.9750
7.0767
7.1613
COS
0.0003
0.9900
0.8237
0.68/7
0.5712
0.4749
0.3952
Cu
O.OOJU
0.9i>OU
0.9553
J . 0 .' j I
O.dOUO
64
6J.5JOO
0.9492
0.9461
0.9430
64.2312
64.2^7
64.2697
64.2^37
L-65
-------
TEMP.
EOUIL. K
FORWARD K
BACKWARD K
•••T IM£»*»
EOUIL.
O.COOO
0.1000
0.2000
0.3000
0.4000
0.5000
0.6000
0.7000
o.eooo
0.9000
1.0000
1. 1000
0.55000E 03 DEO. F
0.80129E 04 0.11728E 04 0.11322E 11 0
0.60000E 05 0.67000E 03 0.78000E 03 J
0.33497E 04 0.57126E 00 0.37130E-01 0
S02
0.7265
2.9900
1.6402
1.5247
1.4754
1.4383
1.4078
1.3824
1.3611
1.3434
1.3284
1.3159
1.3053
H2S
1.1381
3.1900
0.6266
0.5267
0.5387
0.5569
0.5734
0.5877
0.6001
0.6108
0.6200
0.6279
0.6347
SB
0.6665
0.0000
0.5154
0.5592
0.5777
0.5917
0.6031
0.6125
0.6204
0.6269
0.6323
0.6369
0.6407
n20
23.1965
20.0000
22.7628
22.8798
22.8752
22.8625
22.8506
22.8400
22.3308
22.8228
22.8159
22.8099
22.8048
.13032L 19
.65000E 02
.29930E-09
C02
8.9444
7.070v
7.2908
7.4293
7.5429
7.6377
7.7173
7.7841
7.8403
7.8876
7.9276
7.9615
7.9901
CJS
0.0u03
0.95JJ
O.S01 1
0.6710
0.5623
0.4716
0.3958
U.3324
0.2794
0.2350
0.1979
0.1667
0.1405
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j.OOJ^
•J • 7 c J J
0.7636
0.7317
0.7797
0.777&
0.7756
0.7736
0.7716
0.7697
0.7o77
0.7658
0.7639
J • . W V I
63.2/71
6'..77<:3
64.0197
64.3400
64.65^4
64.86dl
0-..901/;
6-.9063
6«4.91 10
1-66
-------
TEMP. o.
EOUIL. K 0.
FORWARD ic o*
BACKWARD K 0.
•««T IME»««
EOUIL.
c.oooo
0.1000
0.2000
0.3000
0.4000
0.5000
0.6000
0.7000
0.8000
0.9000
1.0000
1.1000
1.2000
1.3000
1.4000
1.5000
1.6000
1.7000
1.8000
1.9000
2.0000
2.1000
2.2000
2.3000
2*4000
2.5000
2.6000
2.7000
2.8000
2.9000
3.0000
3.1000
3.2000
TOL. ON EOUIL.
0.17B66E-24
0.17823E-24
55000E
80104E
60000E
38380E
502
1.1258
3.6000
2.1317
2.0364
1.9869
1.9472
1.9144
1.8872
1.8646
1.8456
1.8298
1.8164
1.8052
1.7956
1.7875
1.7806
1.7746
1.7694
1.7650
1.7610
1.7576
1.7545
1.7518
1.7493
1.7471
1.7450
1.7431
1.7413
1.7396
1.7381
1.7366
1.7351
1.7338
1.7324
03 DEC. F
04 0.11728E 04 C.11019E 11 0
05 0.67000E 03 0.78000E 03 0
04 0.57123E 00 0.42539E-01 0
H2S
0.9839
3.3300
0.5259
0.4723
0.4875
0.5031
0.5167
0.5283
0.5382
0.5466
0.5538
0.5599
0.5652
0.5696
0.5735
0.5767
0.5795
0.5820
0.5841
0.5859
0.5875
0.5890
0.5902
0.5914
0.5924
0.5933
0.5942
0.5950
0.5957
0.5964
0.5971
0.5977
0.5963
0.5988
S8
0.7383
0.0000
0.5627
0.5988
0.6174
0.6322
0.6443
0.6544
0.6627
0.6695
0.6752
0.6800
0.6839
0.6872
0.6900
0.6923
0.6942
0.6958
0.6971
0.6983
0.6993
0.7001
0.7008
0.7014
0.7019
0.7023
0.7027
0.7031
0.7034
0.7036
0.7039
0.7041
0.7043
0.7044
H20
23.5753
20.0000
23.0232
23.0911
23.0834
23.0738
23.0652
23.0578
23.0513
23.0458
23.0411
23.0370
23.0336
23.0306
23.0281
23.0259
23.0241
23.0225
23.0212
23.0200'
23.0190
23.0161
23.0173
23.0166
23.0160
23.0155
23.0150
23.0146
23.0142
23.0138
23.0135
23.0131
23.0128
23.0125
.13043E 19
.65000E 02
.37036E-39
C02
8.1742
6.2000
6.4269
6.5701
6.6876
6.7852
6.8664
6.9342
6.9910
7.0386
7.0786
7.1123
7. l«.0d
7.1650
7.1856
7.2032
7.2164
7.2315
7.2429
7.2529
7.2617
7.2693
7.2764
7.2827
7.2384
7.2936
7.2985
7.3030
7.307*
7.3111
7.3149
7.3185
7.3220
7.3253
COS
0.0002
0.9500
0.7945
0.6593
0.5476
0.4552
0.3788
0.3155
0.2631
0.2196
0.1834
0.13J4
0. 12o5
0.1078
0.0906
0.0763
U.Oo*'.
0.0545
0.0462
0.0394
0.0336
0.02o9
0.0249
0.0216
0.0139
0.0166
0.0147
0.0131
0.0118
0.0106
0.0097
0.0089
0.0083
0.0078
Co n2
O.OOCO O.Cuol
0.9800 0.30CO
0.8839
0.8808
0.8776
0.8745
0.8713
0.8682
0.865 1
0.86/1
0.8591
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0 . a 5 3 i
0.8501
0.8471
0.8442
0.6413
0.8J84
0.8355
0.8326
0.8298
0.8269
0.8241
0.8212
0.81b4
0.8156
0.81^8
0.81U1
0.8073
O.B045
0.8018
0. 7991
0.7963
0. '936
CONSTANT NOT MET
0.
0.
25532E 15
25563E 15
0.21649E
0.21020E
03 0.65435E 12
03 0.61762E it
0.27234E
0.272J3t
13 0.22099E *L
13 0.220*8t 'b7
64.2400
64.0028
64. 9034
64.9200
64.9337
64.9452
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64 '.96^o
64,9oVo
64.9/33
64 . 984*.
64. 9d 79
64.9910
64.9962
63.0000
65.0017
63.0032
65.0045
65.0050
63 . OOo?
03.0060
65.0090
63.0100
63.0109
63.0115
65.01 i1*
63.014^
63.0130
1-67
-------
TEMP. 0.70000E 03 DEG. F
EOUIL. K 0.47915E 03 C.44B23E 03 0.96269t 06
FORWARD 1C 0.82000E 05 O.llOOOti 04 O.lJ^O'Ot 03
BACKWARD < 0.42089E 05 0.24540E 01 0.11-olt C2
0. 34 19d
EOUIL.
S02
1.6032
H2S
2.7570
C.4Q02
" 4- • J \J ^ W
0.0000
0. 1000
0.2000
0.3000
0.4000
0.5000
0.6000
0.7000
3.2300
2.2606
2.1992
2.1662
2.1468
2.1341
2.1249
2.1176
3.1700
1.7794
1.9323
2.0127
2.0564
2.0816
2.0972
2.1079
o.ooco
0.3683
0.3894
0.3998
0.4050
0.4077
0.4091
0.4099
20.JOJ.J
21.534*.
21.3911
21.3162
21.2738
21.2331
21.2393
21.2332
6.1200
6.7266
7.0061
7.15U7
7.236 /
7.290*
7.3249
- 7.3509
1.07j,'
0.3292
0.27^0
0.1436
0.07ob
0.0433
0.02d2
0.0194
J.93^>.
0.96-3
0.9439
0.92 /o
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0.6733
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5 3 . 0 3 O i
03 . J / j3
63.03^0
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63 • 0 '»'« '
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1-68
-------
TEMP. 0.70000E 03 u>EG. F
EOUIL. 1C 0.47870E 03 0.44823E 03 0.96177E 08 0.34208E 15
FORWARD K 0.82000E 05 0.11000E 04 O.lOOCOt 05 0.30000E 03
BACKWARD K 0.44430E 05 0.24540E Ol 0.12?88E 02 0.25297t-06
EOUIL.
0.0000
0.2500
0.5000
0.7500
1.0000
1.2500
1.5000
1.7500
2*0000
2.2500
502
1.8688
3.4800
2.4109
2.3736
2.3562
2.3427
2.3305
2.3190
2.3082
2.2981
2.2885
H2S
2.6106
3.3800
1.9535
2.0271
2.0497
2.0633
2.0749
2.0857
2.0958
2.1055
2.1146
58
0.4117
0.0000
0.4039
0.4130
0.4147
'0.4152
C.4155
0.4158
0.4160
0.4162
0.4164
21.6685
20.COOL)
21.5875
21.5207
21.5019
21.4915
21.4823
21.4747
21.4671
21.4599
21.4530
CU2
7.3887
5.6300
6.4050
6.5559
6.6148
6.6566
6.6939
6.7286
6.7613
6.7023
6.8215
COS
0.0019
0.8200
0.1366
0.0302
C.0123
0.0090
O.OCd2
O.OC78
O.OC74
O.OC7J
0.0069
O.OOOJ
O.CJ500
0.3006
0.7263
0.6888
0.6S33
0.61V8
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0.5301
63 .0<«CO
63 . 4nol
03.5071
63.3177
63 . bi6<»
63.3J-«3
63 . S^C
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63.333>
63.301i
1-69
-------
TEMP. 0.70000E 03 OEO. F
EOUIL. K 0.47915E 03 0.44823E 03 0.96269E J?
'FORWARD < o*8230oE oe> ;*iioo-ot j4 c.uru-t J3
BACKWARD K 0.42089E 05 3.2454QE 01 0 c
IME***
EOUIL.
0.0000
0.1000
0.2000
0.3000
0.4000
0.5000
0.6000
0.7000
0.8000
0.9000
.1.0000
1.1000
1.2003
1.3000
1.4000
1.5000
1.6000
1.7000
1.8000
1.9000
2.0000
2.1000
2.2000
2.3000
2.4000
2.5000
2.6000
2.7000
2.8000
2.9000
3.0000
3.1000
S02
1.6032
3.2300
2.2606
2.1992
2.1662
2.1468
2.1341
2.1249
2.1176
2.1112
2.1054
2.1000
2.0947
2.0897
2.0848
2.0800
2.0753
2*0708
2*0663
2.0619
2.0576
2.0534'
2.0493
2.0453
2*0414
2.0375
2.3338
2.0301
2.0265
2.0229
2*0195
2.0161
2.0127
H2S
2.7570
3. 1700
1.7794
1*9323
2.0127
. 2.0564
2.0816
2.0972
2.1079
2.1160
2.1227
2.1286
2.1341
2.1393
2.144-
2.1493
2.1540
2.1587
2.1633
2.1678
2.1722
2.1765
2.1807
2.1849
2.1890
2.1930
2.1969
2.2037
2.2045
2.2082
2.2119
2.2154
2*2189
se
0.4302
0* 0.500
0.3683
0.3894
0.3998
0.4050
0.4077
0.4091
0.4099
0.4103
0.4105
0.4137
3.4108
0.4109
3.4109
3.4110
0.4111
0.4111
0.4112
0.4112
0.4112
0.4113
0.411 i
0*41 1*
0.4U-
0.41 1<>
0.4115
C.4H5
0.4115
0.4115
0.4116
0.4116
0.4116
r* J
21.3133
2.1. JJ .I-.-
21.3344
21.3911
21.3162
21.2738
21.2531
21.2393
21.2302
21.2235
21.2132
21.^1 J3
21.2092
21.2052
21.2013
21.197o
21.193V
21.1903
21.1868
21.1833
21.179*
21.1766
21.1734
21.1702
21.167C
21.16J9
21.1609
21.157V
21.1530
21.1321
21.1493
21.1466
21.1433
O • £ ' 1 i
6.1 i e )tj
• i C */ tr
6.72ob
7 • 0061
7.1347
7.2387
7.290*
7.324*
7.330-y
7. J7,;j
7.3912
7.4J53
7. -2 30
7.w<..,6
/•••36i
7.-710
7. 465o
7.4VV3
7.5iio
7.527-
7. 31,^0
7.3339
7.5667
7. 57* 3
7.3916
7.bj J7
7.0133
7.62 7i
7.6353
7.o4Vb
7.6603
7.6712
7.681V
1 . 3 7 J .
0.2720
G . i <• 3 3
0.07a6
0.043 J
0.02o2
-J.0i7«4
3.01-7
0.01* J
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O.OlJi
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o.ou»<»
U.Ow^i
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O.OC37
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0.0073
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0.9096
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03.oJ i J
03.0304
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03 . 1313
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03.133C
O3.1=0J
o7
1-70
-------
2. INTERMEDIATE REACTOR STUDIES
-------
TABLE OF CONTENTS
2.1 INTRODUCTION
2.2 SUMMARY
2.3 CONCLUSIONS AND RECOMMENDATIONS
2.4 EXPERIMENTATION AND DISCUSSION
2.5 REFERENCE
2-1
2-1
2-2
2-2
2-7
EXHIBIT NO. 2-1
EXHIBIT NO. 2-2
EXHIBIT NO. 2-3
EXHIBIT NO. 2-4
EXHIBIT NO. 2-5
EXHIBIT NO. 2-6
EXHIBIT NO. 2-7
EXHIBIT NO. 2-8
EXHIBIT NO. 2-9
EXHIBIT NO. 2-10
EXHIBIT NO. 2-11
EXHIBIT NO. 2-12
EXHIBIT NO. 2-13
EXHIBIT NO. 2-14
EXHIBIT NO. 2-15
EXHIBIT NO. 2-16
Theoretical Feed Composition
Initial Phase Experimental Data .
Exit Cone, vs Contact Time Plots
at 932°F
7o Conversion - Contact Time Data
7> Conversion vs_ Contact Time Plots
Final Phase Experimental Data
Computer Equilibrium Data
7o Conversion - Temperature Data
70 Conversion vs Temperature Plots
at 1/2 sec.
7o Conversion vs Temperature Plots
at 1 sec.
7o Conversion vs Temperature Plots
at 3/2 sec.
Comparison: Equi. and Exptal. Data
Optimum Conditions for ^9570 Conversion-
Temperature Profiles
Equipment Set-Up Diagram
Sampling System Diagram
2-8
2-9
2-10
2-11
2-12
2-13
2-14
2-15
2-16
2-17
2-18
2-19
2-20
2-21
2-22
2-23
2-i
-------
2.1 INTRODUCTION
2.1.1 In the high temperature reduction of S02 with methane,
there are considerable concentrations of by-products
formed in the primary reactor. Along with some unreacted
S02) Sx, and H2S, there are present tt?, CO, COS and CS2
constituents which represent yield losses and except for COS,
remain unreacted in the subsequent Glaus operation. Since
high temperature primary reactor products normally contain
negligible CS2 concentration, only COS, CO and H2 were
considered for investigation.
2.1.2 Without a reactor -- intermediate between the primary
methane reductor and the Glaus units -- that could
react out these by-products with S02, reductant efficiency
and sulfur yield would suffer. The function of the
intermediate catalytic reactor or converter therefore, is
to insure that sulfur, H2S, S02 and inerts are the only
species entering the Glaus units. This report summarizes
the work performed aimed at optimizing the intermediate
reactor conditions of temperature and contact time (or
space velocity) while attempting to effect maximum
conversion or reduction of COS, CO and H2 to minimal con-
centrations.
2.2 SUMMARY
2.2.1 A laboratory investigation was conducted to establish
the optimum temperature and contact time for the
intermediate reactor. The study was subdivided into two
phases: (a) investigation of variable contact time from
1/4-4 seconds using a fixed temperature of 500°C(932°F); and
(b) study of various temperatures from 350-600°C(662-1112°F) using
the optimum range of contact time obtained. These
experimental parameters were correlated accordingly
thereafter based on the conversion or consumption of COS,
CO and H2 components. Contact times used in this report
are obtained by using a 50% voidage for the catalyst used.
2.2.2 Parameters fixed in the study were feed composition
and type of catalyst. The intermediate reactor feed
used was made of a synthetic gas mixture simulating a typical
product of a high temperature primary SOs/methane reactor.
Porocel LPD was used as catalyst contained in a tubular
Vycor glass reactor.
2.2.3 Complementary computer-based equilibrium calculations
were conducted for the nominal feed to verify the
actual experimental results.
2-1
-------
2.3 CONCLUSIONS AND RECOMMENDATIONS
2.3.1 Initial conclusion derived from the first phase of the
study indicated that at 932 °F, contact times of about
1/2 second for COS, 1 second for CO and 5 seconds for H- were
sufficient to reduce these gases to minimum concentrations.
The optimum contact time range obtained was 1.0 1" 0.5 second.
2.3.2 At 1112°F and 1 second contact time, COS, CO and H-
were essentially reacted out. The general trend
insofar as the intermediate reactor conditions of temperature
and contact time are concerned is that optimum conversion was
favored at higher temperatures and shorter contact times or
lower temperatures and longer contact times.
2.3.3 Each individual component in question exhibited its own
preferred optimum temperature and contact time. Of the
three, Ms was the most difficult to react out followed by
CO, then COS. H2 was completely consumed at 1112°F and 1 second
or at 1022°F and 3/2 seconds. CO was removed at 1058°F and
1/2 second or at 1022°F and 1 second. COS required 752°F and
1 second or 662°F and 3/2 seconds.
2.3.4 Results of this study, and in collaboration with data
gathered from the overall reaction runs of the parallel
normal temperature Claus (NTC) studies conducted at relatively
lower temperature range of about 400-700°F, indicated a
necessity for an intermediate reactor. With the three subject
by-products combined, only COS reacted out favorably in the
NTC 700°F overall reaction runs and confirmed by our 662°F and
752°F runs. CO and H2 needed temperatures higher than the NTC
temperature range to be completely reacted.
2.3.5 Comparison between experimental results and computer
generated equilibrium compositions indicated very close
agreement as the experimental conditions approached the optimum.
2.4 EXPERIMENTATION AND DISCUSSION
2.4.1 Apparatus
An existing test stand or apparatus was modified to
conform with the requirements of the intermediate reactor
studies. An illustrative diagram of the equipment set-up
is shown in Exhibit NOo 2-15. The set-up was composed mainly
of a gas feed manifold or flowmeters control panel, nitrogen
HaO saturator, tubular glass reactor enclosed in a three-zone
vertically mounted furnace and stack. Exhibit No. 2-16 is a
diagram of the sampling system used for both the feed and exit
gases.
2-2
-------
2.4.2 Catalyst
The catalyst used in this study was 1/4" x 3/8" Porocel
LPD. The material was supplied by Porocel Corporation of
Little Rock, Arkansas. A 10-inch high catalyst bed was used and
charged into a 2-inch diameter tubular Vycor glass reactor.
After every week of usage during the experimentation, the
spent catalyst is replaced with fresh batch in order to insure
uniformity of results. A 50% voidage figure was used in
calculating the actual contact times used in this report.
2.4.3 Theoretical Feed Composition
2.4.3.1 As aforementioned, the intermediate reactor
feed composition used was typical of a high
temperature S02/methane reduction product. The first feed
composition proposed for this study was to contain, volumewise:
2.5% S02, 2.0% H2S, 1.0% COS, 1.0% CO, 1.0% H2, 5.0% C02,
11.0% H20 and 76.5% N2. COS, CO and H2 were fixed at 1.0%
each in order to have a common basis for later comparison of
performance. However, one important component absent in the
above composition was sulfur. Since it was obviously
difficult to meter sulfur as it is, concentrations of H2S and
S02 were modified from 2.0% and 2.5% to 5.0% and 4.0%,
respectively, the other components remain unchanged. The
main reason for increasing H2S and S02 concentrations was to
allow sufficient Sx formation }.n the feed stream in order to
better simulate the actual primary methane reductor product
gas composition which contain Sx. Another aim was to provide
H2S and S02 concentrations in the intermediate reactor
exit stream which are suitable for feed to the Glaus converters.
2.4.3.2 The modified feed composition is found in
Exhibit No. 2-1. CS2 and unreacted CEt
components were not included in the feed composition because
they occur in negligible concentrations in most high
temperature S02/CH4 reduction processes.
2.4.3,3 The concentrations of the reacting constituents
in the feed are balanced in proper
stoichiometric proportions such that the following reactions
prevail:
2 COS + S02 > 2 C02 + 3/x Sx (1)
2 H2S + S02 > 2 H20 + 3/x Sx (Claus) (2)
2 H2 + S02 > 2 H20 + 1/x Sx (3)
2 CO + S02 > 2 C02 + 1/x Sx (4)
For every 2 moles each of COS, CO, H2 and H2S there should be
present one equivalent mole of oxidant S02. Therefore, for
the combined 8 moles of the reductants in the feed 4 moles
of S02 would be needed.
2-3
-------
2.4.3.4 Other reactions are also possible, some
occurring to a very limited extent if at all:
3 S2 = Se (5)
4 S2 - S8 (6)
3 H2 + S02 =• H2S + 2 H20 (7)
COS + H20 = C02 + H2S (8)
2.4.4 Laboratory Procedure
2.4.4.1 The laboratory procedure followed for
completing a run consisted mainly of the
following steps: (a) taking the temperature profile;
(b) sampling; (c) GC analysis; and (d) evaluation of results.
2.4.4.2 Before a run was started, the whole system
was first brought under steady state conditions
by feeding into the reactor all the necessary gas flows
through the flowmeters control panel (see Exhibit No. 2-15)
and adjusting the heat input through the three-zone temperature
controller until the catalyst bed temperature becomes constant.
An isothermal condition was established by frequently checking
the temperature profile of the catalyst bed along the 10-inch
bed height at 2" intervals and the midpoint at the 5" level.
A profile reading of the isothermal bed was recorded for every
run made.
2.4.4.3 After taking the temperature profile, three
feed and three exit samples for each run were
then obtained for subsequent gas chromatographic analysis.
Anhydrous CaCl2 was used as the drying agent to remove H2S
and elemental sulfur from the samples (see Exhibit No. 2rl6).
2.4.4.4 In performing the analysis, the GC instrument
was first run with 3 to 5 shots of dilute H2S
gas to condition the fractionating column, as this gas is the
most troublesome to analyze. Actual samples were then analyzed
after the instrument was conditioned. Analytical results were
subsequently evaluated to correlate the pertinent parameters
involved.
2.4.5 Discussion of Results
2.4.5.1 Experimental conditions and analytical results
of all runs are summarized in the exhibits
that follow. Exit analysis is presented as "normalized" values
as derived from actual chromatographic analysis. Due to the
resulting volume decrease from removal of H20 and Sx in the
exit stream, all the remaining components increase correspondingly
2-4
-------
thus bringing forth an inherent error in the actual GC
analysis. Calculations for normalized exit values are
based on maintaining a similar N2 concentration in both
feed and exit streams, the other components being adjusted
accordingly. This was done since N2 was the only component
in the feed whose concentration did not change basically in
the exit. All the others including H20 and C02 were subject
to change.
2.4.5.2 Experimental data of the initial phase of
investigation are shown in Exhibit No. 2-2.
Normalized COS, CO and H2 exit concentrations are plotted
against contact time at a constant temperature of 932*?
in Exhibit No. 2-3 to show the rate of disappearance of
said components relative to contact time. Percent
conversions or consumptions of these components are
tabulated in Exhibit No. 2-4 and Exhibit No. 2-5 plots
these conversions versus contact time.
2.4.5.3 Initial phase results indicate that for
optimum conversions of COS, CO and H2 at
932°F, it takes about 1/2 second for COS, about 1 second
for CO and about 5 seconds for H2 to reach equilibrium
concentrations. 1.0 T 0.5 second was the optimum contact
range obtained.
2.4.5.4 The original plan for the final phase of
study'to test variable temperatures from
662-1292°F.in 90°F intervals at contact times of 1/2, 1 and
3/2 seconds was modified when results obtained at 1112°F
indicated that higher temperature runs were not needed.
Maximum conversion or reduction of COS, CO and H2 to minimal
concentrations has been demonstrated at about 1022-1112°F as
shown by the experimental data in Exhibit No. 2-6.
2.4.5.5 Computer-based equilibrium calculations for
the theoretical feed are given in Exhibit
No. 2-7 at temperatures of 662-1112°F in 90°F intervals.
Actual experimental results are compatible with this data
as temperatures are increased to bring contact times into
practical ranges.
2.4.5.6 Percent conversions of COS, CO and H2 at the
aforesaid ranges of pertinent parameters are
are tabulated in Exhibit No. 2-8. Graphical correlation
between percent conversion, contact time and temperature are
plotted in Exhibit Nos. 2-9, 2-10, and 2-11 corresponding
to contact times of 1/2, 1 and 3/2 seconds, respectively.
2.4.5.7 From the computer calculated equilibrium
compositions of Exhibit No. 2-7, COS, CO
and H2 percent conversions were calculated and compared
with actual experimental results in Exhibit No. 2-12.
2-5
-------
Experimental conversion values were arbitrarily taken from
the 3/2 second contact time results where the components
in question had sufficient data that approached equilibrium
conditions at some particular temperatures. The dotted
lines in this exhibit which correspond to certain temperatures
indicate that experimental values below these lines have
attained equilibrium conditions at the indicated temperatures
and above.
2.4.5.8 For COS, conversion trend is identical in both
experimental and equilibrium values of
Exhibit No. 2-12 although the quantities have very slight
discrepancies. In CO, however, there is very close agreement.
The discrepancy for }{2 is somewhat appreciable due to the
fact that the very slight inaccuracy in GC analysis is
magnified in the extremely low concentration range of H^,.
This occurred even with a highly sensitive GC such as the
Perkin-Elmer 820 we used which contained a very responsive
hot wire detector. Other types of GC with less sensitive
thermal conductivity detectors can hardly identify H;5 at
about 1.0 volume percent range and lower. Despite the
above, however, the overall agreement is quite evident
as the experimental data approached equilibrium conditions.
2.4.5.9 Exhibit No. 2-13 shows the conditions of
temperature and contact time required to
obtain 95% or better conversion for COS, CO and H2. If
US is present in substantial amount in the primary reactor
product, the intermediate reactor conditions should be at
1112°F and 1 second contact time to completely consume this
by-product. At such conditions, however, COS and CO have
already been eliminated. It is apparent therefore, that
higher temperature and longer contact time favor Hs removal
whereas it takes much lower temperatures and shorter contact
times for both COS and CO to disappear. Hence, in a
situation where either COS or CO is abundant in the primary
reactor product and the other by-products are negiligible
the intermediate reactor could operate as low as 662 F
or 932°F, as the case maybe.
2.4.5.10 For an isothermal catalyst bed, the
tolerated variance in temperature during
the experimentation was about i"9 °F although in most cases
fluctuation of about l4°E was common. Typical temperature
profiles of the 662-1112'°F runs along the 10" catalyst bed
height are shown in Exhibit No. 2-14.
2-6
-------
2.5 REFERENCE
Fleming, E. P. and Fitt, T. C., "High Purity Sulfur from
Smelter Gases - Reduction with Natural Gas", Industrial
and Engineering Chemistry, Vol. 42, No. 11, November 1950,
pp. 2249-2253.
2-7
-------
EXHIBIT NO. 2-1
Intermediate Reactor
Theoretical Feed Composition
(Volume %)
Component
C02
COS
H2S
CS,o
S02
H2
0-
N.-:
CH-i
CO
H20
Total
Wet Basis
5.0
1.0
5.0
-
4.0
. 1.0
-
72.0
-
1.0
11.0
100.0
Dry Basis
5.6
1.1
5.6
-
4.5
1.1
-
81.0
-
1.1
-
100.0
2-8
-------
EXHIBIT NO. 2-2
Initial Phase Experimental Data
Run
No.
IR-3
IR-1
IR-5
IR-2
IR-6
IR-4
IR-7
Contact
Time,
(Sec.)
0.345
0.69
1.0
1.38
2.0
2.76
4.0
Nominal
Temp.
CF)
932
932
932
932
932
932
932
Gas Composition (Volume 70)
Feed (Dry Basis)
S02
3.78
4.12
3.51
3.58
4.14
4.35
5.22
H2S
5.18
4.57
4.09
4.57
4.61
5.60
6.36
CO
1.20
1.14
0.92
1.43
1.45
1.55
0.54
COS
1.60
1.01
0.87
0.81
0.84
0.75
1:87
}{•=>
0.79
0.80
0.68
0.83
0.87
0.90
1.44
C02
4.27
6.01
6.14
5.88
6.08
5.68
9.76
N2
83.17
82.28
83.73
82.83
81.90
81.15
74.79
Normalized
Exit (Dry and S-Free Basis)
SO*
1.64
1.58
1.33
1.64
2.01
1.32
1.11
H2S
3.10
1.93
2.27
2.87
3.68
3.20
4.40
CO
0.47
0.10
0.03
0.04
0.01
0.0
0.01
COS
0.04
0.03
0.03
0.02
0.02
0.02
0.06
H2 : co2
0.53
0.43
0.29
0.32
0.20
0.22
0.14
6.25
7.90
7.68
8.00
8.95
7.15
10.42
Np
83.17
82.28
83.73
82.83
81.90
81.15
74.79
Average Feed Concentrations: (Excluding IR-7)
CO = 1.28%
COS = 0.98%
H2 = 0.81%
-------
I
I—'
o
EXHIBIT NO. 2jv3
COS. CO & H= Exit
Coocentrations vs. Contact
Time Plots at 932 °F
-:;3rr. .-. : ::_F"~~T~..:TTTT~" l^f
M Z '^
rr~rrri-i"T~ ~T'" i.~i."i ^"T~.IT."""_.;
".__"! ~.^ : ;i. „„ i _"_• _ _
-:: .•; :-:;;- f :.r I :.':
o
CONTACT TIME, seconds
-------
EXHIBIT NO. 2-4
Contact Time Optimization at 932°F
Percent Conversion Table
Run No.
IR-3
IR-1
IR-5
IR-2
IR-6
IR-4
IR-7
ConLact
Time
Sec.
0.345
0.69
1.0
1.38
2.0
2.76
4.0
Nominal
Temp.
•F
932
932
932
932
932
932
932
COS
97.5
97.0
96.6
97.5
97.6
97.3
96.8
CO
60.8
91.2
96.8
97.2
99.3
100.0
98.2
H2
32.9
46.2
57.4
61.4
77.0
75.5
90.3
2-11
-------
loo
& H^ vs. Contact:
Time at 932°F Plots
II I! IK ! i III I II i
CONTACT TIME, Sec.
ntimimunmmmm
O_1 9
-------
EXHIBIT NCL,2_-6
Final Phase Experimental Data
Run
No.
IR-16
IR-15
IR-17
IR-10
IR- 9
IR-11
IR-19
IR-18
IR-20
IR-24
IR- 5
IR-25
IR-22
IR-21
IR-23
IR-13
IR-12
IR-14
Nominal
Temp.
°F
662
662
662
752
752
752
842
842
842
932
932
932
1022
1022
1022
1112
1112
1112
Contact
Time,
Sec.
0.5
1.0
1.5
0.5
1.0
1.5
0.5
1.5
1.5
0.5
1.0
1.5
0.5
1.0
1.5
0.5
1.0
1.5
Gas Composition^ Volume %
Feed
(Dry Basis)
COS
1.05
1.19
1.45
1.22
1.36
1.21
1.24
1.25
1.60
1.38
1.01
1.38
1.55
1.25
1.30
1.39
1.30
1.73
.CO
1.41
0.73
0.73
0.83
0.71
0.81
0.67
0.73
0.84
0.84
1.14
0.79
0.82
0.71
0.67
0.74
0.62
0.78
H^
0.88
0.84
0.71
0.83
0.73
1.39
0.90
0.77
1.06
0.74
0.80
0.79
0.80
0.66
0.59
0.78
0.71
0.72
H2S
5.47
5.58
5.59
5.34
5.22
5.01
5.19
5.60
5.86
5.73
4.09
4.72
5.70
5.54
4.03
5.07
5.46
4.97
SO?
4.69
4.74
4.88
4.33
3.95
4.52
4.35
4.85
4.71
4.63
3.51
4.28
4.62
4.90
3.97
4.46
5.20
4.72
C02
5.70
6.33
5.91
6.56
6.97
7.21
5.95
6.63
6.88
6.21
6.14
5.06
6.39
6.75
6.42
6.35
7.42
7.60
N2
80.77
80.59
80.69
80.85
80.86
79.80
81.71
80.13
79.0
80.44
83.73
82.72
80.03
79.98
82.75
81.20
79.27
79.46
Normalized Exit
(Dry & S-Free Basis)
COS
0.42
0.05
0.01
0.10
0.01
0.01
0.03
0.02
0.03
0.03
0.03
0.02
0.03
0.03
0.04
0.05
0.05
0.05
CO
1.28
0.59
0.58
0.69
0.42
0.46
0.48
0.23
0.09
0.18
0.03
0.01
0.02
0
0
0
0
0
tiv
0.84
0.74
0.60
0.72
0.48
1.01
0.69
0.41
0.55
0.46
0.29
0.26
0.37
0.08
0
0.07
0
0
H2S
0.95
0.98
1.01
1.72
1.45
1.28
2.70
2.41
2.85
3.19
2.27
2.96
3.72
3.26
3.18
3.50
3.00
3.31
S02
1.86
1.96
1.26
1.76
1.38
1.53
2.13
2.09
1.63
2.10
1.33
1.81
2.08
2.51
2.10
1.99
2.33
1.80
CO-r
6.26
6.62
7.55
7.95
8.61
8.61
7.72
8.23
8.42
8.09
7.68
7.23
8.47
8.57
9.42
9.08
9.73
9.54
Ns
80.77
SO. 59
80.69
80.85
80.86
79.80
81.71
80.13
79.0
80.44
83.73
82.72
80.03
79.98
82.75
81.20
79.27
79.46
NJ
I-"
U)
-------
EXHIBIT NO. 2-7
Computer-Based Equilibrium Compositions
Equilibrium Moles
Component
CH4
S02
H20
H2S
C02
CO
S2
H2
COS
CS2
S6
SB
N2
Total
Initial
Moles
0.0
4.00000
11.00000
5.00000
5.00000
1.00000
0.0
1.00000
1.00000
0.0
0.0
0.0
72.00000
100.00000
660 °F
349 WC
0.00000
1.00669
14.98829
2.01103
6.99832
0.00001
0.09391
0.00068
0.00166
0.00000
0.04257
0.81717
72.00000
97.96034
759 °F
404"C
0.00000
1.56003
13.88483
3.11313
6.99510
0.00009
0.41755
0.00204
0.00481
0.00000
0.05749
0.51775
72.00000
98.55282
840°F
449 "C
0.00000
1.96478
13.08012
3.91555
6.99032
0.00030
1.02090
0.00433
0.00938
0.00000
0.04762
0.22285
72.00000
99.25615
939°F
504"C
0.00000
2.10146
12.81361
4.17620
6.98347
0.00111
1.75615
0.01019
0.01541
0.00001
0.01081
0.01622
72.00000
99.88465
1020 °F
549 HC
0.00000
1.99438
13.03315
3.94785
6.97810
0.00278
2.00983
0.01901
0.01911
0.00001
0.00185
0.00098
72.00000
100.00705
m?^
604 "C
- o.ooooo
1.85474
13.32186
3.64010
6.96868
0.00761
2.23986
0.03805
0.02370
0.00002
0.00023
0.00004
72.00000
100.09487
NJ
I
I-1
-------
EXHIBIT NO. 2-8
Conversion Table for COS, CO and Hp
Nominal
Temp. , °F
662
662
• 662
752
752
752
842
842
842
932
932
932
1022
1022
1022
1112
1112
1112
Component
COS
CO
H2-
COS
CO
H2
COS
CO
H2
COS
CO
H2
COS
CO
H2
COS
CO
H2
Contact Time, Sec.
1/2
60 . 0%
9 . 2%
4.6%
91.8%
16.9%
13.3%
97.6%
27.9%
23.3%
97.7%*
78.0%*
40.0%*
98.1%
97.6%
53.8%
96.4%
100 %
91.0%
1
95.8%
19 . 2%
11.9%
99.2%
40.8%
34.2%
98.3%
68.5%
46.8%
96.6%
96.8%
57.4%
9 7 . 6%
100 %
87.9%
96.2%
100 %
100 %
3/2
99 . 3%
20 . 6%
15.5%
99 . 2%
43.2%
27.4%
98 . 2%
89 . 2%
48 . 0%
97.5%*
99.0%*
70.0%*
96.9%
100 %
100 %
97.1%
100 %
100 %
* Interpolated values from Exhibit No. 2-5
of initial phase study. Analytical
results of confirmatory runs for these
interpolated values are shown in
Exhibit No. 6.
2-15
-------
0 .
-------
N
in
10 ,
it
z
Ul
o
u
I
§
M
CO
O
e-s
EXHIBIT NO. 2-10
Conversion vs.
Temperature Plata
I second Contact
TEMPERATURE,
6*0
-------
o
M
3
EXHIBIT NO. 2-L
Conversion vs.
Temperature 'lots
at 3/2 second Contact
Time
TEMPERATURE, "F
-------
EXHIBIT NO. 2-12
Comparison of Percent Conversions Between Experimental
Results* and Computer Calculated Equilibrium Values
Temp. , °F
660
759
840
939
1020
1119
Percent Conversion
COS
Experimental
Results
99.3
99.2
98.2
97.5
96.9
97.1
Equilibrium
Values
99.8
99.5
99.1
98.5
98.1
97.6
CO
Experimental
Results
20.6
43.2
89.2
99.0
~ 100
~ 100
Equilibrium
Values
100
100
100
99.9
99.7
99.2
H2
Experimental
Results
15.5
27.4
48.0
ZQ40 . .
~ 100
~ 100
Equilibrium
Values
99.9
99.8
99.6
99.0
98.1
96.2
N>
* Experimental values were taken from 3/2 second contact time results.
-------
EXHIBIT NO. 2-13
Optimum Conditions Required for
>95% Conversion of COS, CO & Ho
Component
COS:
CO:
H2:
Temp. , °F
662
842
932
1022
1022
1112
Contact
Timej Sec.
1
1/2
1
1/2
.3/2
1
2-20
-------
EXHIBIT NO. 2-14
Typical Catalyst Bed
Temperature Profiled
Nominal
Temp.
°F
662
752
842
932
1022
1112
Actual Bed Temperature, WF
Bed Height from Bottom of Bed
0"
662
759
838
928
1020
1125
2"
658
761
838
928
1016
1119
4"
658
756
842
928
1018
1112
5"
662
752
842
932
1018
1112
6"
666
756
851
932
1022
1112
8"
671
759
849
936
1026
1112
10"
667
757
842
932
1022
1112
2-21
-------
Heating Tape
Burners
N>
Exit
Sample
Point
Exit
Thermocouples
for
Bed Temp.
STACK
— Thermometer
LJL i
Immersion
Heater
Feed
Sample
Point
,H
REACTOR
rhermocouples <; \
for 3-Zone'
Furnace
Temp.
THREE-ZONE
FURNACE
Chimney
Stopcock
to
a
P
N5
The rm-0-Watch
NITROGEN H20
SATURATOR
r .- r r s ' r ' r r ' ' r r '
Insulation
Catalyst
Heating Element
Heating Tape
Insulation
FLOWMETERS
CONTROL
PANEL
COS CO
INTERMEDIATE REACTOR
EQUIPMENT SET-UP~
-------
EXHIBIT NO. 2-16
Main Process Line
Heating
Tape
Exit
SAMPLE
BULB
FILTERING
FLASK
DRYING TUBE
High Pressur
Rubber Tubing
CaCl2
INTERMEDIATE REACTOR
SAMPLING SYSTEM
2-23
-------
3. LOW TEMPERATURE GLAUS STUDIES
-------
TABLE OF CONTENTS
3.1 INTRODUCTION 3-1
3.2 SUMMARY 3-3
3.3 CONCLUSIONS ^__ _ _ _ 3-4
3.3.1 Initiation Temperature 3-4
3.3.2 Heat Responses _^ 3-4
3.3.3 H2S/S02 Reactions 3-5
3.3.4 Effect of NOX _ 3-5
3.3.5 Catalyst Life __ 3-5
3.4 RECOMMENDATIONS 3-5
3.5 DATA AND DISCUSSION OF EXPERIMENTAL PROGRAM __ 3-5
3.5.1 Laboratory Apparatus Used 3-5
3.5.2 Catalysts Used 3-6
3.5.3 Operational Procedures 3-7
3.5.4 Advantages and Disadvantages of Apparatus 3-8
3.5.5 Comparison with Previously Reported Data 3-8
3.5.6 Initiation Temperature 3-9
3.5.7 Heat Responses 3-10
3.5.8 H2S/02 Reactions _ ___ 3-13
3.5.9 Effect of NO* 3-15
3.5.10 Catalyst Life _._ 3-15
3.6 REFERENCES _ __ 3-15
EXHIBIT NO. 3-1-1 Typical Low Temperature Glaus
Process Flow Sheet 3-17
EXHIBIT NO. 3-1-2 Process Profile of Exhibit 3.1.1 3-18
EXHIBIT NO. 3-2 Apparatus Used 3-19
EXHIBIT NO. 3-3 Sampling Apparatus 3-20
EXHIBIT NO. 3-4 Tabulation Demonstrating H2S/S02
Initiation Temperature 3-21
EXHIBIT NO. 3-5 Tabulation Demonstrating the Need
for Staging 3-21
EXHIBIT NO. 3-6 Tabulation Demonstrating the Effect
of Oxygen Content on temperature
Excursions 3-22
EXHIBIT NO. 3-7 Temperature Responses through the
Catalyst Bed with Time 3-23
3-i
-------
TABLE OF CONTENTS
(Continued)
Page
EXHIBIT NO. 3-8 Tabulation Demonstrating the Heat
Induced by Water Adsorption 3-24
EXHIBIT NO. 3-9 Tabulation Demonstrating the Effect
of Water Presaturation of the
Catalyst 3-24
EXHIBIT NO. 3-10 Comparison of 100% and 97% Conversion
in the Claus Reactor 3-25
EXHIBIT NO. 3-11 Experimental Determination of Sulfur
Yield Loss --- 3-26
EXHIBIT NO. 3-12 Tabulation Demonstrating H2S Reactions
Over Activated Alumina 3-27
EXHIBIT NO. 3-13 Comparison of Average Analysis of
Sulfur Bearing Streams 3-28
EXHIBIT NO. 3-14 Tabulation Demonstrating the Effect
of NOX 3-29
EXHIBIT NO. 3-15 Tabulation Demonstrating Catalyst
Degradation 3-29
APPENDIX 3-1 Calculation of Adiabatic Temperature
Rise in the H2S/S02 Reaction 3-30
APPENDIX 3-2 Compilation of Experimental Runs in
Low Temperature Claus Program 3-32
3-ii
-------
3.1 INTRODUCTION
3.1.1 The classic Glaus Process consists of burning one
third of an H2S stream completely to sulfur dioxide
in a waste heat boiler and reducing this with the remaining
two thirds H2S over a catalyst, usually activated bauxite,
at 400-650°F. The two stages of H2S removal can be
represented in the following way.
H2S + 3/2 02 > S02 + H20 (1)
2 H2S + S02 > 3 S + 2 H20 (2)
At these temperatures, sulfur is passed through the Glaus
reactor in the vapor phase and subsequently recovered in
a sulfur condenser.
3.1.2 The low temperature Glaus (LTC) process utilizes reaction
(2) to reduce S02 contained in stack gas to sulfur.
The LTC process is distinguished from the conventional Glaus
by operation at a temperature low enough to inhibit oxygen
reactions. This temperature was visualized to be below
400°F. At these temperatures the sulfur produced is assumed
to be retained on the catalyst as solid or liquid sulfur.
Preliminary calculations had indicated that a serious sulfur
loss could occur if the sulfur on the catalyst was exhibiting
its normal vapor pressure of the temperature range of the
LTC process. This means that if a large loss as sulfur vapor
is to be avoided, the catalyst must act as a sulfur vapor
adsorbent as well as a catalyst. This is one assumption
which had to be validated in the LTC laboratory program.
Previous works had suggested that catalysts such as activated
alumina were strong sulfur adsorbents. Information available
at the time of conceptual process design also indicated that
20yo sulfur by weight can be deposited on these catalysts before
significant impairment of catalytic activity occurs. In the
LTC process the sulfur, thus deposited, is then removed and the
catalyst regenerated by heating to about 900°F and purging with
an inert or reducing gas. The sulfur is recovered by cooling
the purge stream and condensing the sulfur. One third of the
sulfur is taken as product and the remaining two thirds is
used to synthesize the H2S required for the reduction process.
The regenerated catalyst is recycled to process. H2S can be
synthesized with sulfur and H2, or with sulfur, steam and C1U.
When CH4 is used as the reactant, the H2S synthesis gases will
contain small amounts of COS and CS2. The H2S efficiency in
the LTC unit would be adversely effected if the COS and CS2
did not react to the same extent as H2S reacts. The extent
of COS and CS2 reaction is another aspect of process design
3-1
-------
which would have to be investigated in the LTC laboratory
program. Because 2/3 of the sulfur is recycled a high sulfur
yield must be obtained within the LTC unit. A 97% yield
from the Claus unit produces only a net sulfur yield of 9170
from the entire system. The low temperatures of the LTC
process do however favor a high equilibrium conversion
and high sulfur yields were anticipated.
3.1.3 Princeton Chemical Research Incorporated, previously
contracted (Contract Number PH 86-68-48) to perform
similar work with low concentration powerhouse stack gases.
They conducted an extensive catalyst screening program from
which we selected the prime catalyst candidates for our work.
3.1.4 Allied Chemical's principle concern with the LTC
process was its possible applicability to the
smelter situation, specifically copper smelting.
A typical copper smelter's stack analysis was taken as
follows:
Component Volume %
S02 2.9
02 14.3
CO 0.6
C02 1.7
H20 0.1
N2 80.4
100.0
Some 76% of S02 emissions of primary non-ferrous smelters
are from copper smelters. Current public and governmental
concern is high in the S02 pollution area and S02 emissions
are increasing at an alarming rate. This points out the
growing urgency for a useful S02 abatement process.
In Phase I of Contract PH 22-68-24 Allied Chemical reviewed
the technology of S02 abatement by reduction to
sulfur techniques- At that time the low temperature Claus
process appeared to be the most promising approach to the
direct reduction of S02 in smelter gas.
3.1.5 Exhibits 3.1-1 and 3.1-2 illustrate one version of
how LTC was visualized to be implemented for copper
smelter stack gases. Exhibit 3.1-1 shows one modification
of this process applied to our weakest smelter gas with
2.9 percent S02 and 14.3 percent oxygen. In this case,
H2S is synthesized from sulfur and reformed methane in the
equipment shown in the lower left of Exhibit 3.1-1. Half
of the H2S is mixed with the smelter gas, and the mixture
is then cooled to about 150 °F by direct water injection prior
3-2
-------
to entering the first of two stages in a moving catalyst bed
type of reactor. The adiabatic heat rise of about 112°F
necessitates further intercooling before the gas, mixed with
the remaining half of the H2S, enters the second bed. Gases
venting the reactor are incinerated and stacked.
Catalyst discharging from the reactor, containing about
20 percent sulfur by weight, is fed continuously to the regen-
erator, which operates at about 930°F. This is a multiple
tube type heat exchanger with the catalyst inside the tubes.
Heat is supplied by boiling sulfur and condensing the vapors
on the outside of the tubes. An inert or reducing gas sweep
is recycled through the regenerating catalyst mass and the
sulfur recovery condenser.
Catalyst leaving the regenerator is cooled by an air
sweep prior to re-entering the reactor. Somewhat over two
thirds of the sulfur is recycled for H2S synthesis, and the
remainder taken as product.
3.2 SUMMARY
3.2.1 The objective of the low temperature Claus program was
to (1) verify assumptions and optimize conditions on a
laboratory scale using a simulated smelter gas feed and (2) confirm
the laboratory results on an actual smelter gas using similar,
but portable, equipment. The LTC concept is based on reducing S02
with H2S in the presence of 02 at a temperature where 02 will not
react in the system. To develop a usable process several
aspects of process design had to be investigated and
established. First the temperature at which the LTC reaction
initiates had to be ascertained. The lowest operating tempera-
ture limit was set at the dew point of the gas mixture which
was anticipated as feed for the Claus unit. The lower
temperature limit for 02 reactions also had to be established.
The heat responses caused by water adsorption on the catalyst
and its effect on reaction and reaction temperature was
another important design consideration. If these heats
imposed a significant temperature increase within the
catalyst bed interferring 02 reactions could ensue. The
presence of NOx in smelter gas prompted an investigation
on its effect on catalyst activity and its influence on
subsequent yield. Activated alumina and molecular sieve
were the prime catalyst candidates for the LTC process and
their effectiveness and life characteristics had to be
determined. The catalyst used must retain a high degree
of catalytic activity. Successful LTC operation is based
on no less than 97% conversion. Below this the process is
inoperable from both an air pollution and economic standpoint.
3-3
-------
3.2.2 A laboratory program was initiated to establish the
salient aspects which comprise the design of a LTC
process. One unique advantage of our laboratory work stemmed
from design and use of a reactor system which behaved
substantially adiabatically. Heretofore laboratory
investigators utilized reactors which performed under apparent
isothermal conditions. The adiabatic nature of our reactor
system permitted the observation of several previously
unobserved heat responses. Upon further investigation the
LTC process was found unworkable because of the resultant
myriad of complications. High temperature responses
accompanied by interferring 02 reactions, a detrimental
contribution from NOx and subsequent loss of catalytic
activity resulted in sulfur yield losses that rendered the
LTC approach untenable. Consequently, the field tests were
not carried out.
3.3 CONCLUSIONS
3.3.1 Initiation Temperature
Low temperatures are reported to inhibit 02 reactions
and favor sulfur yield in the LTC environment. It was felt
previously that the high heat of reaction involved in the
H2S/S02 reaction would cause the reaction to go to completion
in the LTC operating range once it was initiated. 120 F was
the lowest temperature considered because it approached the
dew point of a visualized smelter gas feed stream (cooled
by direct water injection). Fresh activated alumina was
expected to be an active catalytic media over which reaction
would initiate at a low temperature but with repeated catalyst
use the initiation temperature was anticipated to rise. This
change in initiation temperature was not experienced as the
catalyst consistently initiated reaction at 120°F. Although
only a limited time was spent with molecular sieve, the same
consistency in reaction initiation at 120°F was experienced.
Consequently we can conclude that all the catalysts studied,
1/4-1/2" Alcoa F-l, 1/4" Alcoa H-151 and 1/16" Linde molecular
sieve 13X, will initiate the Glaus reaction at 120°F.
3.3.2 Heat Responses
The adiabatic temperature rise due to H2S/S02 reaction
in a typical air-diluted smelter gas is approximately 82°F/
7o S02 converted. This excludes other heat inducing variables
such as water and sulfur adsorption on the catalyst. The
actual temperature rise experienced was consistantly higher
than anticipated. These higher than anticipated temperatures
increase the probability of 02 reactions, decrease the
equilibrium conversion and increase the loss of sulfur values
in the vapor phase. The net result of this inordinate tempera-
ture rise was an unacceptable sulfur yield loss which by
itself makes LTC an inoperable process.
3-4
-------
3.3.3 HgS/Qg Reactions
One of the original premises upon which LTC was based
was that H2S would not react with 02 at LTC temperatures.
This was proven false by passing H2S and 02 through alumina
catalyst. These were strong indications of 02 reaction at
temperatures as low as room temperature. Subsequent analyses
of inlet and exit streams did show H2S oxidation does occur
in the LTC range.
3.3.4 The Effect of NOX
The presence of 100 ppm NOX consistantly had a
detrimental effect on sulfur yield. This was true with S02
concentrations in both the smelter and power stack gas range.
Thus NOx increased the already unacceptable yield loss due
to temperature rises.
3.3.5 Catalyst Life
With continued use the catalytic activity of both
Alcoa F-l and H-151 decreased.. After each run the catalyst
was regenerated at 950°F by passing first a N2 purge through
the catalyst bed and then a N2 purge stream containing
50% H2S. Catalyst activity loss is attributed to sulfate
formation on the catalyst. H2S in the purge stream was
intended to remove the sulfate by reduction to sulfur.
This method of regeneration proved ineffective.
3.4 RECOMMENDATIONS
Data developed in this study show the LTC process
for direct reduction of S02 in stack gases to be inoperative.
It is recommended that no further work be done on this
approach.
3.5 DATA AND DISCUSSION OF EXPERIMENTAL PROGRAM
3.5.1 Laboratory Apparatus Used
3.5.1.1 The reactor system (see Exhibit 3-2) consisted
of a three inch diameter flanged stainless
steel tube which contained the catalyst bed supported by a
1/4" mesh stainless steel grid. Four five inch 2100 watt
band heaters, individually controlled with rheostats,
furnished heat to the reactor. The temperature was monitored
by four iron-constantan thermocouples which were strategically
placed within an axial thermowell. The temperature was
recorded on a Honeywell temperature recorder. The system
3-5
-------
was well Insulated and behaved substantially adiabatically
at the low temperatures which were utilized. Gases were
fed across the top of the reactor bed, passed through the
catalyst bed, through the reactor's lower leg and subsequently
through a combustion tube. The lower leg, fabricated from
Vycor, was heated with heating tapes. This provided visual
inspection for traces of elemental sulfur which escaped the
catalyst bed. The temperature in the lower leg was monitored
with an iron-constantan thermocouple and recorded on the
Honeywell temperature recorder. The combustion tube was
two inches in diameter and contained an eight inch bed of
Alcoa F-l activated alumina. It was heated with a twelve
inch 3000 watt band heater and its shell temperature was also
monitored. The combustion tube was maintained at 950 °F
to oxidize any sulfur gases which passed through the reactor.
This served as a second monitor for any sulfur which might
have broken through the reactor system, as gas analysis at
this point would show an excess of SOs relative to the gases
monitored directly after the reactor. The feed gases were
metered by Fisher- Porter tri-flat rotameters. Air and
nitrogen streams were passed through a humidifier to establish
the necessary water concentration just prior to mixing with
the sulfur bearing gases. Heating tapes provided heat to the
lines between the humidifier and the reactor system. This
prevented the possibility of water condensation in the lines.
Three sample points were included, one' at the inlet to the
reactor tube, one at the reactor tube exit and one at the
exit of the combustion tube. Samples were taken by passing
the gas stream at the sample point through a drying media and
through a 250 ml gas sampling bulb. The samples were then
analyzed on a Perkin-Elmer Gas Chromatograph (Model 810) for
H2S, S02, N2 and Q2»
3.5.2 Catalysts Used
3.5.2.1 Three different catalysts were used in the
low temperature Glaus program^
(1) 1/4 inch Alcoa H-151 Activated Alumina.
(2) 1/4 - 1/2 inch Alcoa F-l Activated Alumina.
(3) 1/16 inch Linde 13X Molecular Sieve.
These catalysts were choosen on the basis of previously
reported work by others. The bulk of the work was done
with the activated aluminas because of their comparative
low cost. Linde molecular sieve underwent only a short
test period because it induced higher temperatures and
consequent lower sulfur yields than the aluminas. Both
aluminas were almost equivalent in effectiveness although
the H-151 consistently exhibited a more favorable temperature
response.
3-6
-------
3.5.3 Operational Procedure
Run Procedure
3.5.3.1 The catalyst bed temperature was stabilized
by passing nitrogen, equivalent in volume
to the run flow volume, through it and adjusting the
band heaters to enforce the required temperature profile.
This was performed with bone dry nitrogen except when the
catalyst was preloaded with water.
3.5.3.2 After the catalyst bed had been stabilized,
the nitrogen flow was replaced with the feed
flow. Water was added to the feed stream by passing the
nitrogen and air through a constant temperature water bath
to produce the required water concentration. The lower
reactor leg was maintained at 400°F to eliminate the
possibility of sulfur buildup if the sulfur was not
efficiently adsorbed in the catalyst bed. The combustion
tube was maintained at 950°F to insure that any sulfur
bearing gases would oxidize and exit as S02. This served
as a check against the possibility of elemental sulfur
leaving the reactor system undetected. The gas chromato-
graph could easily analyze S02 whereas sulfur would have
been condensed out in the sampling system and gone
unnoticed.
3.5.3.3 Samples were taken and analyzed on the
Perkin-Elmer 810 gas chromatograph periodically.
Regeneration Procedure
3.5.3.4 At the end of each run, if the catalyst was
to be reused, it was regenerated. The
temperature of the catalyst bed was regulated to 900-950°F
by adjustment of the band heaters. A dry nitrogen purge
stream of 300-500 cc/min was passed through the catalyst
bed overnight (approximately 16 hours). A purge stream
of 500-800 cc/min containing approximately 50% H2S was then
passed through the catalyst for one half hour. During the
entire regeneration procedure excess air was passed through
the lower reactor leg and combustion tube to eliminate the
possibility of molten sulfur buildup. The catalyst bed
temperature was then reduced by passing approximately
60 liters per minute of dry nitrogen across the catalyst bed.
It took several hours to cool the entire catalyst bed to
operating conditions for the next run.
Sampling Procedure
3.5.3.5 Exhibit 3-3 illustrates the sampling apparatus
used. At the start of each run a fresh supply
of CaCl2, the drying agent used, was charged to the drying tube.
3-7
-------
A 250 ml open ended gas sample bottle was connected to the
end of the sampling line to collect a sample and the
sampling valve was opened. The process gas was passed
through the sample system for a 5 minute period prior to
closing and removing the sample bulb. This guarded against
the possibility of adsorption losses in the drying agent
and insured a valid homogeneous sample in the sample bulb.
The drying tube charge was changed after forty minutes of
use. This prevented the CaCl2 from becoming spent.
3.5.4 Advantages and Disadvantages of Apparatus
3.5.4.1 One of the outstanding advantages of the
equipment was the adiabatic nature of the
reactor. This permitted the observance of heat responses
which had been previously overlooked with smaller and
poorly insulated reactors. Smaller equipment has a tendency
to operate isothermally. Once the catalyst bed temperature
had been stabilized it remained constant until influenced
by internal reaction and adsorption phenomenon. The
temperature was monitored and recorded continuously at
three points in the catalyst bed. Another advantage was
the adaptability to variance in feed conditions which the
system had. It operated with S02 feed concentrations varying
from 0.3 tc 3.0 percent.
3.5.4.2 The major disadvantage of the apparatus
was that its size and adiabatic nature
made temperature control difficult. It took time for the
band heaters to heat the three inch diameter bed uniformly.
Also the catalyst bed's adiabatic nature caused cooling to
take an even longer time. Because of this slow temperature
response to coding runs could not always be performed as
often as desired.
3.5.5 Comparisons with Previously Reported Data
Princeton Chemical Research, Inc.
3.5.5.1 Princeton Chemical Research was contracted
(Contract Number PH 86-68-48) by the
Department of Health, Education and Welfare for the
Development of Processes to Reduce Sulfur Dioxide to
Elemental Sulfur.PCR became involved in low temperature
Claus as it applied to a typical power stack gas. Power
stack gas generally contains approximately 0.3% S02 by
volume as compared to approximately 3.0% contained in
air-diluted mixture of reverberatory and converter exits
in a copper smelter.
3-8
-------
3.5.5.2 Several catalyst were screened by PCR.
Linde molecular sieve 13X, Alcoa F-l
alumina and Alcoa H-151 alumina were found to be among
the best overall performers. It Is these catalysts which
were used in Allied's LTC program.
3.5.5.3 PCR performed their experimentation in a
one inch reactor with a 1/4" axial thermowell
as compared with the three inch reactor which Allied used.
PCR used only about 15 grams of catalyst while Allied used
approximately 1700 grams. The considerably smaller size
of the PCR apparatus gave them the capability of running
their experimentation at a rapid pace. Allied felt a
larger reactor would be advantageous in obtaining realistic
data. The larger equipment effectively operated adiabatically,
as previously reported, and it is primarily this feature
that allowed measurement of heat responses, heretofore
unobserved, that resulted in concluding the LTC process
unacceptable. It should be noted, however, that heat
release due only to H2S + S02 reaction in the 0.3% S02
range is small and would be inconsequential if water, oxygen,
and NOX were not present.
Consolidation Coal Company
3.5.5.4 Consolidation Coal presented a paper at the
September 1969 meeting of the American
Chemical Society entitled "Removal of Sulfur Dioxide from
Power Plant Stacks by a Modified Glaus Process". Their
work parallels that of Princeton Chemical Research, Inc.
3.5.5.5 The catalysts which they used were quite
similar to catalysts used both by Allied
and PCR - high surface area activated aluminas.
3.5.5.6 Consolidation Coal used a reactor tube
34 mm (1 1/3 inches) in diameter with an
axial thermowell. The catalyst bed was one to three inches
in height as compared with Allied's 3 inch diameter reactor
tube with a 15 inch catalyst bed. Consequently the apparatus
used is much more similar in size and application to that
of PCR and has the same disadvantages and advantages
associated with it.
3.5.6 Initiation Temperature
3.5.6.1 Exhibit 3-4 demonstrates that the H2S-S02
reaction does initiate over alumina catalyst
as low as 120°F. Runs 11, 14 and 17 all resulted in an
appreciable temperature response. Exit analysis also
indicated a considerable reduction in both H2S and S02.
3-9
-------
3.5.7 Heat Responses
3.5.7.1 The temperature rise in the catalyst bed was
not anticipated to exceed the adiabatic
temperature rise due to reaction alone. The adiabatic
temperature rise, approximately 82°F/% S02 reacted was
calculated as illustrated in Appendix 3-1.
3.5.7.2 In the smelter gas range the temperature rise
in a single bed is excessive and staging
with intercooling on a commercial unit would be required.
The adiabatic temperature rise in a smelter gas containing
370 S02 would be 3 X 82°F or 246°F. This temperature rise
would have an adverse effect on equilibrium and most probably
induce oxygen reactions. Exhibit 3-5 illustrates this
staging requirement. In run 4 a 6.0% H2S - 3.0% S02 stream
was fed over F-l alumina at 260°F in the presence of 02.
Within five minutes the temperature had already risen to
650°F. At this point the feeds were stopped to curtail
reaction. On a commercial scale this catastrophic tempera-
ture rise would prove damaging not only in terms of conversion
but also in terms of equipment. Runs 5 and 7 repeat this
6.0% H2S - 3.0% S02 feed without 02. The temperature rise
is prohibitive as far as yield is concerned but certainly
not the drastic temperature rise experienced with 02. In
run 4 we were obviously experiencing oxygen reactions.
The temperature change in run 7 is greater than that
experienced in run 5. This may be due to the fact that
there was more sulfur condensation and adsorption at the
lower initial bed temperature, 150°F as compared with 260°F.
3.5.7.3 Exhibit 3-6 demonstrates that 02 does effect
temperature excursions in the Glaus
environment. In runs 20-24 approximately 15% 02 replaces
N2 28 minutes into the run. In each instance there is an
immediate temperature response. This implies that oxygen
reactions are proceeding.
3.5.7.4 Noteworthy is the fact that the adiabatic
temperature rise as calculated is not inclusvie
in that it does not account for temperature rises due to
adsorption phenomenon, especially water and sulfur.
R. E. Derr of Alcoa in an article in Industrial and Engineering
Chemistry in April 1938 discussed the use of activated alumina
as a drying agent. He states:
In the adsorption of 1 pound of moisture from
any gas, heat approximately equivalent to the
condensation of one pound of steam is liberated.
If air containing 7 grains of moisture per
cubic foot is being dried, then about 1000 BTU
3-10
-------
will be converted for every 1000 cubic foot
of air dried. This is sufficient to raise
the temperature of air, at normal pressure,
about 30°C (54°F). One might expect that
the temperature of the dry exit gas would
immediately show a corresponding rise, but
this is not the case. The alumina stores
the heat in the zone where adsorption at
high efficiency is taking place, and the
temperature of the exit gas remains below
the calculated figure throughout more than
half the adsorption period. The temperature
of the exit then rises above the mean
temperature, but adsorption at high efficiency
continues until the exit gas approaches its
maximum temperature.
3.5.7.5 As Derr indicates water is a significant
heat contributor and we can infer that
this heat would be realized in the operation of the low
temperature Glaus unit. Indeed the heat responses that
were typical in our low temperature Glaus system passed
through the reactor in the same manner as Derr explained.
Exhibit 3-7 illustrates the typical heat response exhibited
across the low temperature Glaus unit with time. This type
of heat response was exhibited whenever there was heat
liberated as a result of passing gas through the catalyst
bed. Exhibit 3-8 demonstrates that the heat liberated due
to water adsorption is significant. Runs 1-3 utilize
Alcoa F-l alumina (1/4-1/2"). As the length of the run
was extended more of the catalyst bed exhibited a heat
response. The catalyst used in run 1 was dehydrated by
passing a N2 purge through it at 300°F for 1/2 hour.
This proved to be an ineffective method for dehydration
as run 2 demonstrated (the top of the catalyst bed saw
only a 95°F temperature rise as compared with 150°F in
run 1). The catalyst used in run 2 was then dehydrated
at 500°F for 1/2 hour. The temperature response at the
top of catalyst bed in run 3, 125°F, showed that 500°F
is much more effective in terms of dehydration. Runs 8
and 9 utilized Alcoa H-151 (1/4"). The contact time was
decreased from 7.1 seconds (superficial at S.T.P.) to
2.0 seconds. As might be expected a decrease in contact
time was reciprocated with an increase in the rate of
temperature response. One can draw the conclusion that
the rate of temperature response is inversely proportional
to the contact time. Because the H-151 and F-l aluminas
are so similar in nature, this increase in the rate of
temperature response is not due to differences in the two
aluminas. In Run 9 13% or approximately four times the water
concentration as in run 8 was fed into the reactor. An
interesting observation is that although both catalyst
bed attain approximately the same maximum temperature,
3-11
-------
run 8 achieves it approximately four times faster. It may
be concluded from this that the rate of temperature response
is directly proportional to inlet water concentration under
these operating conditions.
3.5.7.6 The recognition that the heat of water
adsorption on alumina catalyst is appreciable
prompted the following speculation:
Use of dry, regenerated catalyst can
give a heat rise which will be the sum of the
heat of reaction and the heat of adsorption
of water. This can be true even if dry gases
are reacted, since water is formed in the
reaction. This additive effect probably can
be avoided if the catalyst, after regeneration
is loaded with water and cooled prior to the
reaction. Again, if a cooled water-loaded
catalyst is used, it is just possible that
the sulfur being retained in the catalyst will
tend to displace the adsorbed water. Since
heat effects are reversible, this would provide
cooling that would counteract the heat of
reaction. Indeed preliminary calculations
indicate that with a 15 percent water loading on
the alumina, and a subsequent 20 percent loading
with sulfur, the system would be substantially
isothermal regardless of the HaS-SOs concentration.
However, the phenomena of water displacement
by sulfur adsorption must occur, otherwise this
concept will not work.
3.5.7.7 Exhibit 3-9 tabulates our work in this area.
In each instance water-saturated nitrogen
was fed across the catalyst bed until the bed temperature
stabilized at the prescribed initial bed temperature for
that particular run. It generally took several hours for
the catalyst bed temperature to stabilize. In run 6 a
6.0% H2S-3.0% S02 stream was fed across water presaturated
F-l alumina at 150°F. A temperature rise up to 425°F was
experienced. The phenomenon of water displacement by sulfur
adsorption obviously did not induce isothermal conditions
in run 6. In runs 12 and 13 run 5 was repeated with a
3.0% H2S-1.5% S02 feed stream at a lower initial bed
temperature. The thought was that a lower temperature
alumina may retain enough additional water to induce
isothermal conditions. Runs 12 and 13 exhibited very little
temperature change, 35°F maximum, but the exit analysis
indicated very little reaction taking place. Run 14 repeated
runs 12 and 13 but using a dry catalyst bed. This would
ascertain if the catalyst had somehow become poisoned. A
high temperature response and a much higher yield with the
3-12
-------
dry catalyst was noted. Identical results were attained in
runs 15, 16 and 17, showing both low temperature and low yield
responses with a water presaturated catalyst and substantially
the opposite with a dry catalyst. It was concluded from
this series of runs that water presaturation at 120°F
deactivates the catalyst.
3.5.7.8 Exhibit 3-10 demonstrates why high sulfur
yields in the Glaus reactor are important.
Because two thirds of the sulfur produced is recycled to the
reactor system a 97% yield in the Glaus reactor produces a
net yield loss of approximately 9%. Exhibit 3-11 illustrates
how yield loss was experimentally determined. Exit samples
were taken during each run and plotted as pound sulfur per minutes
versus time. The area under this curve was then computed and
compared with the pounds sulfur which would have been collected
on the catalyst at 10070 conversion.
3.5.8 HgS/Og Reactions
3.5.8.1 Several runs were performed to establish the
temperature threshold at which H2S oxidation
occurs. This is important because the LTC process is based
on operating in a temperature regime where reaction interference
from Os is insignificant. Activated alumina catalysts
(F-l and H-151) were utilized in these tests as they appeared
to be the prime candidate for the LTC reaction. Both
regenerated and fresh alumina were tested in the following
manner. H2S and N2 were passed across the alumina, heated by
the reactor band heaters, thereby stabilizing the temperature
of the catalyst bed. 02 in the form of air was then substituted
for a portion of the nitrogen establishing a 15% 02 content
in the feed. Samples were collected and analyzed at the inlet
and exit and the temperature in the catalyst bed was monitored
throughout the run. Exhibit 3-12 tabulates the feed and
resultant response in several runs. Exhibit 3-13 compares
the average analysis of sulfur bearing streams into and out of
the catalyst bed. In each instance an immediate temperature
rise occurred indicating an exothermic reaction was proceeding
on the alumina. In most cases the exit stream contained both
H2S and S02 and the total of the sulfur bearing gases exiting
the bed was not equivalent to the H2S entering the bed.
This implies that the formation of sulfur did take place
across the alumina. The reduction of H2S in the exit stream
can be accounted for by combinations of the following
reactions:
H2S + 1/2 02 > S + H20 (1)
S + 02 > S02 (2)
2 H2S + S02 > 3 S + 2 H20 (3)
H2S + 3/2 02 > S02 + H20 (4)
3-13
-------
3.5.8.2 In runs 28-31 regenerated H-151 alumina was
used. At the inlet temperature of 100°F
in run 29 59% of the H2S passed through the reactor with the
bulk of the difference being converted to sulfur. This was
accompanied by a maximum temperature change of 105°F.
In run 30 with an inlet temperature of 130°F similar results
were attained. Only 1% of the HaS passed through the reactor
unreacted in run 31, which had an inlet temperature of 205°F.
677o went to S02 and the remainder to sulfur. The maximum
temperature change realized was 715°F. Run 28 demonstrates
that at 315PF essentially all of the H2S was converted to
S02.
3.5.8.3 In summation, the following can be deduced
about the net reactions occurring with
regenerated H-151 alumina:
(a) Above 315°F, only reaction (4) takes place.
(b) Above 205°F reaction (4) dominates.
(c) Somewhere below 205°F reaction (1) dominates.
3.5.8.4 In runs 69-70 fresh H-151 alumina was used.
In run 69 the inlet temperature was 75°F
and the maximum temperature change experienced was 825°F.
Exit analysis indicated that 45.5% of-the H2S feed was
converted to S02, 12.5%, remained unreacted and the balance
went to sulfur. At a inlet temperature of 150°F, 69.5%
of the H2S went to S02, 3% did not react and 27.5% was
converted to sulfur. The maximum temperature change in the
catalyst bed was 830°F. The implication here is that fresh
H-151 is more reactive than regenerated H-151. The
predominate net reaction occurring was reaction (4).
3.5.8.5 Fresh F-l alumina was used in runs 65-67.
The inlet catalyst bed temperature in
run 65 was 75°F, in run 66 it was 165°F and in run 67,
250°F. Sulfur was deposited on the catalyst by running a
preliminary short term LTC reaction on the catalyst used in
runs 66 and 67. The presence of sulfur would demonstrate
whether or not sulfur would be stripped from the catalyst
if H2S reactions did occur. In run 65 more than half (53%)
of the H2S passed through the catalyst bed unreacted, the
balance forming sulfur. In both runs 66 and 67 the
major reaction which took place was reaction (4). The
analyses also indicated that reaction(2) also took place to
a significant degree. The temperature response in both
runs 66 and 67 registered off the recorder scale (in excess
of 945°F). From these runs it can be concluded that if
H2S oxidation does occur to an appreciable extent, sulfur
on the catalyst will also b« oxidized. Also, at 75°F H2S
oxidizes predominantly to sulfur and at 165°F and above,
predominantly to S02.
3-14
-------
3.5.9 Effect of
3.5.9.1 Smelter gases contain minute amounts of N0y
and its presence has been found to promote
sulfate formation on alumina catalysts causing a loss of
catalyst activity. Exhibit 3-14 demonstrates that the
presence of only 100 ppm NOx consistently induces a detrimental
effect on LTC yields. Also the smaller the inlet concentration
of H2S and S02 the more significant is the NOx deactivation
phenomenon.
3.5.9.2 0.6% H2S and 0.3% S02 were passed across
H-151 alumina in runs 35, 38 and 45. No
NOx was present and the average yield loss is 12.770.
Identical runs 36, 37 and 46 with NOx present in the feed
resulted in a 26.2% yield loss. With a 1.2% H2S/0.6% S02
feed 25.0% yield loss was realized with NOx a™J 17.0%
without (runs 43 and 44). At the 3.0% H2S-1.5% S02
concentration level (runs 32, 33, 34 and 40) an average
yield loss of 18.4% was experienced without NOX and
26.0% with NOx present.
3.5.10 Catalyst Life
3.5.10.1 To make the LTC operation a viable process
the catalyst used must be capable of being
stripped of sulfur and rejuvenated to its original state
of activity. Following each run the alumina was heated to
900 °F+ and a N2 purge was passed across it to remove the
sulfur. Next a purge stream containing 50% H2S was passed
through the catalyst bed to reduce sulfate formed in the
LTC reaction. This regeneration procedure had reportedly
been used successfully in prior work. This procedure,
however, did not reactivate the catalyst to its previous
state as indicated in Exhibit 3-15.
3.5.10.2 In run 41 fresh H-151 alumina was used.
After 15 hours use, which involved four
regenerations, two duplicate runs were made (runs 45 and 46).
In each instance the yield loss was increased significantly.
A repeat of this procedure was made with F-l alumina. In
run 47 yield loss over fresh F-l was effectively zero.
After ten regenerations and 26 hours use two identical runs
(runs 57 and 58) produced an 8-10% yield loss.
3.6 REFERENCES
1. Belgian Patent 661,381, issued to Peter Spence, LTD,
7/16/65.
3-15
-------
2. British Patent 1,097,306, issued to Peter Spence, LTD,
1/3/68.
3. British Patent 1,132,846, issued to Peter Spence, LTD,
11/6/68.
4. Arthur G. McKee Company. Systems Study for Control of
Emissions. Primary Nonferrous Smelting Industry, for
NAPCA, Contract No. PH 86-65-85, May 1969.
5. R. T. Struck et al, (Consolidation Coal Co.), "Removal of
Sulfur Dioxide by a Modified Claus Process", presented
at ACS meeting of September 1969.
6. Gamson & Elkins, "Sulfur from Hydrogen Sulfide",
Chemical Engineering Progress, Vol. 49, No. 4, April 1953,
pp. 203-215.
7. R. E. Derr, "Drying Air and Commercial Gases with Activated
Alumina", industrial & Engineering Chemistry, April 1938,
pp. 384-8.
3-16
-------
EXHIBIT 3-1-1
HpS REDUCTION OF SOg
LOW TEMPERATURE GLAUS
VENT
HgS COOLER
INTERCOOLER
HoS SYNTHESISX
CO
REFORMER
METHANE
SULFUR
COOLER
®>
SMELTER
SPRAY
COOLERS
GA^
SULFUR
PUMP
SULFUR
BOILER
REGENERATE
CATALYST
COOLER
SULFUR
PRODUCT
RECOVERy
^
&
.VCI E SULFUR
-------
EXHIBIT NO. 3-1-2
lav Temperature Glaus
210000 SCTM 2C Smelter Gas
Process
Points
Temp. °C
SCFM
CFM at T°C
Dew Point °C
% CH4
% EsO
% CO
% H2
% C02
% COS
% HaS
% S2
% S8
% S02
% 02
I N2
A
25
3073
3360
100.00
B
205
9217
16140
100.00
C
898
18320
78570
0.35
29.03
11.57
54.20
4.85
D
300
19850
41690
0.32
26.78
10.68
50.00
4.48
7.74
E
625
19300
63480
0.11
19.03
3.96
19.05
10.53
0.15
35.69
.
F
300
19300
40530
0.11
19.03
'3.95
19.05
iO_.,53
0.15
35.69
K
540
18640
55510
16.74
0.11
1.31
16.17
0.34
64.98
0.34
L
125
18580
27090
16.80
0.11
1.31
16.23
0.34
65.21
0.00
M
227
210000
384720
0.10
0.60
1.70
2.90
14.30
80.40
N
125
9290
13540
16.80
0.11
1.31
16.23
0.34
65.21
0.00
0
223
219290
398450
0.81
0.58
0.05
2.32
0.02
2.76
2.78
13.68
77.00
P
65
243290
301190
47.5
10.60
0.52
0.05
2.09
0.02
2.50
2.50
12.34
69.38
Q
115
234200
332800
11.01
0.54
0.05
2.17
0.00
0.00
1.31
12.82
72.08
R
125
9290
13540
16.80
0.11
1.31
16.23
0.34
65.21
S
116
243490
346980
11.23
0.53
0.10
2.70
0.02
2.50
1.26
12,33
69.34
T
65
252540
312640
53.4
14.40
0.51
0.10
2.61
0.02
2.40
1.21
11.89
66.86
U
115
249500
354540
16.98
0.51
0.10
2.67
0.00
0.024
0.012
12.03
67.67
V
125
1000
1460
100.00
W
420
3160
8020
37.70
30.80
31.50
Point
X
I
J
Y
Sulfur Distribution
From Catalyst
From Gas Cooler
Total S Recovery
Recycle Sulfur
Net S Recovery
N.T./Day
1166.72
8.15
1174.87
789.83
385.04
Conversion
Equilibrium
Used Here
99.637.
99.50%
Sulfur Yield 98.52%
Methane Factor llA9~0 SCF/Ton S
Catalyst Circulation Rate 194.4 Ton/Hr.
Point
G
H
ZA
ZB
ZC
Temp.
°C
300
250
100
1 75
400
3-18
-------
Agitator
Insulation
Heating
Tape
Humidifier
Bypass Valves
Thermostatically
Controlled
Heater
Inlet
Sample
Point
Thermocouples
Catalyst
Therraowell
Exit
1 & Sample
f Point
Thermocouple
Humidifier
5" Band Heat
Insulation
Thermocouples
Exit
Sample
Point
Supporting
Screen
12" Band Heater
1/4" F-l Alumina
Supporting
Screen
Heating
Tape
-------
EXHIBIT 3-3
SAMPLING APPARATUS
Process
Line
u>
K>
O
Sample
Line
** o
Sam
A Valve
Heating
Tape
Glass
Wool
250 Ml
Sample Bulb
CaCl2
Drying Tube
Rubber
Tubing
-------
EXHIBIT 3-4
TABULATION DEMONSTRATING H3S/SOp INITIATION TEMPERATURE
Run
No.
11
14
17
7.
H2S
3.0
3.0
3.0
7.
SQa
1.5
1.5
1.5
%
N2
85.5
85.5
85.5
%
02
None
None
None
%
H20
10.0
10.0
10.0
NOX
(ppm)
None
None
None
Total
Flow
(L/min.)
51.2
51.2
51.2
Initial
Bed
Temp.
CF)
120
120
120
Contact
Time
(sec.)
2.0
2.0
2.0
Duration
of
Run
(min. )
75
165
130
Catalyst
Used
Regen.
H-151
Regen.
H-151
Regen.
H-151
AT
Max.
Top
(°F)
100
80
110
AT . AT
Max.
Middle
(°F)
.. 190
185
220
Max.
Bottom
(°F)
170
190
200
Max.
Temp.
(°F)
.310 . .
310
340
EXHIBIT 3-5
TABULATION DEMONSTRATING THE NEED FOR STAGING
Run
No.
4
5
7
%
Hj-S
6.0
6.0
6.0
%
sop
3.0
3.0
3.0
7.
Np
73.8
83.8
80.4
%
OP
10.0
None
None
7.
H20
7.2
7.2
,0.
NOx
(PPm)
None
None
None
Total
Flow
(L/min.)
51.2
51.2
51.2
Initial
Bed
Temp.
CF)
260
260
150
Contact
Time
(sec.)
2.0
2.0
2.0
Duration
of
Run
(min.)
5
105
110
Catalyst
Used
Dehyd.
Regen.
F-l
Regen.
F-l
AT
Max.
Top
(°F)
390
115
230
AT
Max
Middle
(°F)
215
275
AT
Max.
Bottom
(°F)
116
275
Max.
Temp.
(°F)
650
480
425
3-21
-------
EXHIBIT 3-6
EFFECT OF OXYGEN CONTENT ON TEMPERATURE EXCURSIONS
Run
No.
20
21
22
23
2k
7.
H2S
3.0
3.0
3.0
3.0 .
3.0
7o
S02
1.5
1.5
1.5
1.5
1.5
°k
N2
Bal.
Bal.
Bal.
Bal.
Bal.
*
7o
02
16.0
.0
15.0
15.0
15.0
Z
H20
10.0
10.0
10.0
10.0
10.0
NOx
(ppm)
None
100
None
100
None
Total
Flow
(L/min.)
51.4
51.4
51.4
51.4
51.4
Initial
Bed
Temp.
(°F)
120
120
150
150
170
Contact
Time
(sec.)
2.0
2.0
2.0
2.0
2.0
Duration
of
Run
(min . )
80
75
72
60
75
Catalyst
Used
Regen .
H-151
Regen.
H-151
Regen.
H-151
Regen.
H-151
Regen.
H-151
AT
Max.*
Top
(°F)
10
None
10
20
15
AT
Max.*
Middle
(°F)
40
30
65
135
165
AT
Max.*
Bottom
(°F)
30
50
90
235
265
Max.
Temp.
(°F)
360
370
390
545
585
-AT max. measured from temperature achieved after 28 minutes into run.
At this point 02 is added.
3-22
-------
EXHIBIT 3-7
TEMPERATURE RESPONSES THROUGH THE CATALYST BED WITH TIME
A
to
M
u>
H
ro
B
13
(D
H
(o
n
C
H
Time
-------
EXHIBIT 3-8
TABULATION DEMONSTRATING THE HEAT INDUCED BY WATER ADSORPTION
lun
to.
I
2
3
8
9
%
HaS
None
None
None
None
None
I
SO,
None
None
None
None
None
%
N*
89.4
89.4
89.4
97.0
87.0
%
0=
None
None
None
None
None
%
HoO
10.6
10.6
10.6
3.0
13.0
NO,
/Don)
None
None
None
None
None
Total
Flow
(L/min. )
7.2
7.2
7.2
51.2
51.2
Initial
Bed
Temp.
rn
120
120
120
120
120
Contact
Time
(sec. )
7.1
7.1
7.1
2.0
2.0
Duration
of
Run
fmin.)
60
60
340
100
100
Catalyst
Used
Fresh F-l
Dehyd.F-1
Dehyd.F-1
Fresh H-151
Dehyd.H-151
AT
Max.*
Top
(°F)
150
95
125
35
60
AT
Max *
Middle
(°F)
None
None
115
115
115
AT
Max.*
Bottom
(°F)
None
None
100
100
85
Max.
Temp.
(°F)
270
215
245
235
235
EXHIBIT 3-9
TABUIATION DEMONSTRATING THE EFFECT OF WATER PRESATURATION OF THE CATALYST
lun
*o.
6
12
13
L4
L5
16
17
7.
H2S
6.0
3.0
3.0
3.0
3.0
3.0
3.0
%
S02
3.0
1.5
1.5
>5
1.5
1.5
1.5
%
N2
80.4
85.5
85.5
85.5
85.5
85.5
85.5
%
02
None
None
None
None
None
None
None
1
H20
10.6
10.0
10.0
10.0
10.0
10.0
10.0
NO*
(ppm)
None
None
None
None
None
None
None
Total
Flow
(L/min.)
51.2
51.2
51.2
51.2
51.2
51.2
51.2
Initial
Bed
Temp.
<°F)
150
120
120
120
120
120
120
Contact
Time
(sec.)
2.0
2.0
2.0
2.0
2.0
2.0
2.0
Duration
of
Run
(min.)
125
105
135
165
105
130
130
Catalyst
Used
Regen.H20
Sat. F-l
Regen.H20
Sat. H-151
Reeen.HgO
Sat. H-151
Regen.Dry
- H-151
Regen.H20
Sat. H-151
Regen.HpO
Sat. H-151
Regen.Dry
H-151
AT
Max.*
Top
(CF)
60
5
5
80
5
25
110
AT
Max.*
Middle
(°F)
255
25
25
185
40
35
220
AT
Max.*
Bottom
(°F)
275
30
35
190
85
30
200
Max.
Temp.
(°F)
425
150
155
310
205
150
340 j
3-24
-------
EXHIBIT 3-10
COMPARISON OF 100% AND 97% CONVERSION
IN THE GLAUS REACTOR
Exit Loss as
S02 - 0.0% (Net S02 Conversion-100%)
3% S02 Smelter Gas
6% H2S
H?S
Generator
Low Temp. Glaus
100% Conversion
Recycle Sulfur
Recovered Sulfur
Regen-
erator
66.7% of
Total Sulfur Feed
-» 33.3% oi
Total
Sulfur
Feed
Exit Loss as
S02 - 0.3% (Net S02 Conversion - 91%)
3% S02 Smelter Gas
6% H2S
H2S
Generator
Low Temp. Glaus
97% Conversion
Recycle Sulfur
Regen-
erator
66.7% of
Total Sulfur Feed
Recovered
Sulfur
30.4% of
Total Sulfi
Feed
3-25
-------
EXHIBIT 3-11
EXPERIMENTAL DETERMINATION OF SULFUR YIELD LOSS
POUNDS SULFUR/MINUTE
VS
MINUTES
Pounds Sulfur
Total loss in Exit x 100
Pounds Sulfur Fed
= % Yield Loss
u>
i
N>
a\
Total Yield Loss
in Exit
a
:=>
CO
CO
Q
Data Points Determined
by Exit Analysis
Time in Minutes
-------
EXHIBIT 3-12
TABULATION DEMONSTRATING HgS REACTIONS OVER ACTIVATED ALUMINA
Run
No.
29
30
31
28
69
70
65
66
67
7«
H2S
3.5
3.5
3.0
4.0
4.5
3.5
6.0
6.0
6.0
7.
S02
None
Hone
None
None
None
None
None
None
None
7.
Na
81.5
81.5
82.0
77.0
80.5
81.5
79.0
79.0
79.0
7»
02
15.0
15.0
15.0
19.0
15.0
15.0
15.0
15.0
15.0
7.
H20
None
None
None
None
None
None
None
None
None
NQX
(ppm)
None
None
None
None
None
None
None
None
None
Total
Flow
(L/min.)
51.4
51.4
51.4
51.4
51.4
51.4
51.4
51.4
51.4
in ic Lai
Bed
Temp.
<°F)
100
255
205
315
75
150
75
165
250
Contact
Time
(sec.)
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
2.0
Duration
of
Run
(min.)
63
158
75
145
85
85
85
65
75
Catalyst
Used
Refcen. H-151
Regen.H-151
Regen.H-151
Regen.H-151
Fresh H-151
Fresh H-151
Fresh F-l
Fresh F-l
Fresh F-l
A >• . AT
Max.*
Top
C°F)
45 '
None
120
835
365
85
25
1050+
945+
Max.*
Middle
<°F)
60
10
715
525
800
830
260
585
425
^T
Max.*
Bottom
(°F)
105
40
655
150
825
645
260
600
415
Max.
Temp.
<°F)
290
295
920
1225
900
980
335
1200+
1200+
3-27
-------
EXHIBIT 3-13
COMPARISON OF AVERAGE
ANALYSIS OF SULFUR BEARING STREAMS
Run
29
30
31
28
69
70
65
*66
*67
% HvS in
3.40
3.47
3.32
4.09
4.50
3.80
7.07
6.38
6.73
% HPS out
2.01
2.20
0.03
4.07
0.57
0.12
3.75
0.01
0.06
% SOp out
0.30
0.23
2.23
0.00
2.05
2.64
0.01
14.12
6.84
% S by
Difference
1.09
1.04
1.06
0.02
1.88
1.04
3.31
-
-
* - Sulfur Initially Present on the Catalyst
3-28
-------
EXHIBIT 3-14
TABULATION DEMONSTRATING THE EFFECT OF
Run
No.
35
38
45
36
37
46
43
44
34
40
32
33
I
H2S
0.6
0.6
0.6
0.6
0.6
0.6
1.2
1.2
3.0
3.0
3.0
3.0
^
S02
0.3
0.3
0.3
0.3
0.3
0.3
0.6
0.6
1.5
1.5
1.5
1.5
1
N2
74.1
74.1
74.1
74.1
74.1
74.1
73.2
73.2
70.5
80.5
70.5
70.5
T.
02
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
%
H20
10.0
10.0
10.0
10.0
10.0
10.0
10.0
10.0
10.0
None
10.0
10.0
N°x
(ppm)
None
None
None
100
100
100
None
100
None
None
100
100
Total
Flow
(L/min.)
51.6
51.6
51.6
51.6
51.6
51.6
51.6
51.6
51.6
51.6
51.4
51.4
Initial
Bed
Temp.
(°F)
- 120
150
120
120
120
120
120
120
120
120
120
120
Contact
Time
(sec.)
2
2
2
2
2
2
2
2
2
2
2
2
Duration
of
Run
(min.)
195
375
282
473
102
165
130
130
68
150
103
98
Catalyst
Used
H-151
H-151
H-151
H-151
H-151
H-151
H-151
H-151
H-151
H-151
H-151
H-151
% Yield
Loss
7.0
17.0
14.0
28.0
17.0
33.6
17.0
25.0
19.5
17.2
26.5
25.5
A T
Max.*
Top
<°F)
95
85
100
85
105
95
95
95
115
45
185
215
AT
Max.*
Middle
(°F)
185
200
165
180
160
165
175
160
300
120
520
620
A T
Max.*
Bottom
CF)
185
200
165
175
160
175
175
160
300
125
460
550
Max.
Temp.
<°F)
305
350
285
300
280
295
295
280
420
275
640
740
EXHIBIT 3-15
TABULATION DEMONSTRATING CATALYST DEGRADATION
Run
No,
41
45
46
47
57
58
7.
H2S
0.6
0.6
0.6
0.6
0.6
0.6
7.
SO*
0.3
0.3
0.3
0.3
0.3
0.3
I
No
74.1
74.1
74.1
74.1
74.1
74.1
%
0;.
15.0
15.0
15.0
15.0
15.0
15.0
I
HoO
10.0
10.0
10.0
10.0
10.0
10.0
NOx
foom)
None
None
None
None
None
None
Total
Flow
(L/min.)
51.4
51.4
51.4
51.4
51.4
51.4
Initial
Bed
Temp.
(°F)
120
120
120
120
120
120
Contact
Time
(sec.)
2
2
2
2
2
2
Duration
of
Run
(min . )
200
282
165
98
157
107
Catalyst
Used
H-151
H-151
H-151
F-l
F-l
F-l
7. Yield
Loss
0.7
14.0
33.6
None
10.0
8.3
AT
Max.*
Top
(°F)
120
100
95
45
35
35
AT
Max.*
Middle
(°F)
155
165
165
135
115
35
AT
Max.*
Bottom
(°F)
135
165
175
115
105 k
105
Max
Tern.
(°F
275
285
295
255
235
225
3-29
-------
APPENDIX 3-1
CALCULATION OF ADIABATIC TEMPERATURE RISE
IN THE HgS/SQp REACTION
Basis: 100 moles feed at 125°F (approx. 52°C) with the following
constituent concentrations
H c I ^n This assumes a two stage operation
en * ' i q with the feed simulating the expected
5 " o^ e values in the first stage.
N2 - OO.J
The following reaction takes place:
(1) 3 H2S + 1.5 S02 > 4.5 S8 + 3.0 H20 (100% Conversion
Assumed)
(AHR1)250C - (ZAH°f) prod. - ( 2 A H|) react
- where AH°f - Heat of formation of products and reactants
- 4.5(-.07) + 3.0(-57.80) - 3.0(-4.82)+1.5(-70.95)
- r52.9 k cal.
With the inlet temperature at 125°F and the outlet temperature
unknown:
(AHJ>out -
-------
APPENDIX 3-1
mCp ATin - 725.5 (52-25) - 19500 cal,
mCp of the product (assuming final temp. - 248°F) is
calculated as follows:
Out Moles ^Specific Heat 248°F mCp
H20 13.0 8.10 105.2
Ss 4.5 6.13 27.6
N2 86.5 7.27 627.0
759.8
The change in temperature then is:
AT , 52900 + 19500 _ nc ,. „„
and the final temperature is 95.4°C + 25.0°C - 120.4°C
120.4°C ( « 249°F) matches the assumed 248°F well,
248°F—
therefore, AT/% S02 - . ..' - 82°F/% S02
' - Mean Cp in cal/gm mole/°C
3-31
-------
APPENDIX 3-2
IM TEK'ERATURE CUUS
COMPILATION . OF EXPERIMENTAL DATA
TOP FEED USED
2 SECOND SUPERFICIAL CONTACT TIME A S.T.P.
Comments
Fresh Catalyse
7.1 sec. contact
time for Runs 1,
2, and 3
•Cat Bed Preset with
HS0
Cat Bed Preset with
HzO
Cat Bed Presat with
HuO
Cat Bed Presat with
H20
Cat Bed Presat with
H20
Cat Bed Presat with
H20
Oxygen added \
20 min. Into (
run {
\
Nominal Feed Conditions
Run
I
2
3
I*
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
20
21
22
23
24
25
26
27
28
29
30
31
32
33
34
35
36
I H2S
Hone
None
None
6.0
6.0
6.0
6.0
None
None
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.0
3.33
3.0
3.0
3.0
3.0
3.0
3.0
None
3.0
4.0
3.5
3.5
3.0
3.0
3.0
3.0
0.6
0.6
I S02
None
None
None
3.0
3.0
3.0
3.0
None
None
1.5
1.5
1.5
1.5
1.5
1.5
1.5
1.5
1.5
1.67
1.5
1.5
1.5
1.5
1.5
1.5
None
1.5
None
None
None
None
1.5
1.5
1.5
0.3
0.3
I N2
89.4
89.4
89.4
73.8
83.8
80.4
80.4
97.0
87.0
85.5
85.5
85.5
85.5
85.5
85.5
85.5
85.5
85.5
95.0
70.5
70.5
70.5
70.5
70.5
70.5
85.0
70.5
77.0
81.5
81.5
82.0
70.5
70.5
70.5
74.1
74.1
I 02
None
None
None
10
None
None
None
None
None
None
None
None
None
None
None
None
None
None
None
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
19.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
15.0
">, H-0
10.6
10.6
10.6
7.2
7.2
10.6
10.6
3.0
13.0
10. 0
10.0
10. 0
10.0
10. 0
10.0
10. 0
10.0
10.0
None
10. 0
10.0
10. 0
10. 0
10.0
10.0
None
10.0
None
None
None
None
10.0
10.0
10. 0
10. 0
10.0
NO,
.5
T 0
2e.O
3-32
-------
4. REDUCTION OF STRONG S02 TO SULFUR
-------
TABLE OF CONTENTS
4.1 INTRODUCTION __ __
4.2 SUMMARY
4.3 CONCLUSIONS AND RECOMMENDATIONS
4.4 EXPERIMENTATION AND DISCUSSION
4-1
4-1
4-2
4-2
EXHIBIT NO. 4-1
EXHIBIT NO. 4-2
EXHIBIT NO. 4-3
EXHIBIT NO. 4-4
EXHIBIT NO. 4-5
EXHIBIT NO. 4-6
EXHIBIT NO. 4-7
Equipment Set-Up
Theoretical Feed Compositions
Experimental Data
SOz and CH* Conversions _"
Effect of Temperature __" ..
Effect of Contact Time
Effect of Water
4-6
4-7
4-8
4-9
4-10
4-11
4-12
4-i
-------
4.1 INTRODUCTION
4.1.1 Two major sources of sulfur dioxide contributing to
air pollution are the stack gases from fossil
fuel-fired power generating stations and from primary
non-ferrous smelting operations. The former emits huge
volumes of gas with a low (0.20-0.47o) S02 content. In the
latter area, copper smelters contribute some 77 percent of
the smelter S02 pollution. These gases are variable in both
gas composition and in volume. Both sources are very high
in oxygen.
4.1.2 It is widely recognized that elemental sulfur is the
most desirable final product of S02 pollution abate-
ment processes. This involves reacting the S02 with a
reductant such as methane to eventually yield sulfur.
Neither of the aforementioned gases are suitable for direct
(in-situ) reduction. It has been concluded by Allied Chemical
that the most practical approach to the problem is to use a
gathering process capable of removing the S02 from the stack
gases and delivering to a reduction process a steady flow of
essentially 100% S02 on a dry basis. Process and economics
developed on using, as an example, the DMA absorption-
desorbtion process coupled witn a methane reduction process,
bear out this conclusion. The DMA study is given in
Section 5 of this report.
4.1.3 Accordingly, work was initiated on an experimental
program to define an optimum process profile for
reduction of strong S02 with methane. Due to expiration of
our contract, only a limited amount of data was generated.
The objective of the reported work was to determine the
initiation temperature of the S02-CIU reaction, and the
temperature-contact time required to approach equilibrium.
The primary purpose of determining the initiation temperature
was to define the preheat required for the feed gases to the
reactor.
4.2 SUMMARY
4.2.1 A laboratory investigation was conducted to meet the
aforementioned objectives of the study. Twelve runs
were made for the strong S02 reduction with CK*. The runs
were conducted at conditions of 1000-1500°F at 100°F intervals,
1/2-5 seconds contact time, 0 and 10% HsO in feed and
SOa/CIU feed ratio of 2/1. A proprietary catalyst for high
temperature primary S02 reduction with CH* was used.
4-1
-------
4.2.2 The first part of the study involved investigation of
varying temperatures at contact times of 3 and 5 seconds
to determine the initiation temperature range. In the other
part of the study shorter contact times were investigated at
different temperatures to determine the corresponding
practical contact time. The synthetic S02 feed gas used
simulated a typical product of an S02 gathering plant such
as that using the DMA process. For the purpose of evaluating
the results obtained, calculations for conversions of CH4 and
S02 were based only on the exit GC analysis. Near actual
contact times were used by assuming a 50% voidage for the
catalyst bed.
4.2.3 Preceding this study, theoretical equilibrium calcula-
tions on computer were done for strong S02 reduction
with CH4 by extrapolating from known kinetics developed for
weaker S02 reduction with cm using the same proprietary
catalyst. Actual experimentation subsequently followed to
verify the validity of the extrapolation.
4.3 CONCLUSIONS AND RECOMMENDATIONS
4.3.1 Preliminary results from the initial portion of the
study indicated that at 3 and 5 seconds contact time
the minimum reactor temperature or initiation temperature
range for the strong S02/CH4 reaction was 1400-1500°F. The
initiation temperature at 3 seconds contact time was about
1450°F.
4.3.2 At practical contact times such as 1 or 2 seconds,
1450-1500°F appears to be the optimum initiation
temperature range. For a reactor temperature of 1500°F
and 1 second contact time, conversion of CH4 was about 97%
and that of S02 was about 83%.
4.3.3 The experimental initiation temperatures obtained were
much higher than computer predicted extrapolations
from weaker concentration SOs/CIU kinetics.
4.3.4 Comparison between feeds with 0% H20 and 10% H20
demonstrated no significant difference in both CH4
and S02 conversions.
4.4 EXPERIMENTATION AND DISCUSSION
4.4.1 The same apparatus was used in this investigation as
in the intermediate reactor studies except for some
minor changes such as the use of a different catalyst and
the kind and number of feed gases. A diagram of the equipment
set-up is found in Exhibit No. 4-1.
4-2
-------
4.4.2 The compositions of the feed used in the runs herein
reported contain 0% or 10% H20 on a wet basis.
Introduction of H20 in the gaseous mixture was achieved by
bubbling the CH4 component through a water bath maintained
at the required temperature. Theoretical feed compositions
with 0% and 10% H20 are shown in Exhibit No. 4-2.
4.4.3 Actual gas chromatographic analysis of both feed and
exit samples are found in Exhibit No. 4-3. The feed
analysis is on a dry basis only while the exit is on dry and
sulfur-free basis. The "unnormalized" exit analysis should
not be directly used in conjunction with the feed to calculate
conversions of CH4 and SOa because of volume change due to
reaction and different bases (water free, and water and sulfur
free). By writing equations for C, H, 0 and S balance from
a general SOs/Cm reaction equation it can be shown that the
exit analysis even on water and sulfur free basis is sufficient
to calculate back the actual feed composition (dry basis) and
the various conversions. Hence, for the purpose of evaluating
the results obtained, calculations were based only on the
exit GC analysis. In cases where the calculated inlet feed
composition did not agree well with the intended feed composi-
tion, the run was either repeated or cancelled if there was
an indication of an error in GC analysis or the experimental
set up.
4.4.4 From the exit analysis following the above method,
Cm efficiency, actual SOa/Cm feed ratio used and
conversion of SO 2 to various products have been calculated.
Results of these calculations at certain conditions of
temperature, contact time and H20 content of feed are
tabulated for each particular run in Exhibit No. 4-4.
4.4.5 Our arbitrary definition of minimum reactor temperature
or initiation temperature for the strong SO^Cm
reaction is the temperature at which better than 80% S02
conversion and better than 90% Cm efficiency are achieved
at a contact time of 5 seconds or less.
4.4.6 Based on 100 moles of the exit gas, CH4 efficiency is
calculated by using the following equation:
r,^., , (CH4) in - (CIU) out ._.
Efficiency - J — in - x °
where; (CH4)in • (C02 + CO + Cm + COS + CS2)exit
(cm)out • [(CQ + Hg) + cm] exit
Note: The chemical formulae of the compounds shown in these
equations represent the corresponding number of moles
of the compound.
4-3
-------
The actual SOc/CH4 feed ratio is computed as follows
(S02/CH4)feed ratio -
u'ill(calc)
(SQpUn calc. J 3/2 (CO+COS) + 2 C02 + S02 + CS2 - 1/2 (Hg + H2S)
(CH4)in calc. [(C02 + CO + CH4 + COS + CS2)
4.4.7 S02 conversions to various products are calculated from
the following:
-, ™ „ ., (SQg)in - (SOp)exit v
Total S02 Conversion = — —(S02)in
exit
S02 Conversion to COS = (S0g)in — x 10° = Rl
S02 Conversion to H2S = ^gOali^^ x 10° = R2
S02 Conversion to CS2 = 2((so2Tint x 10° = R3
Unreacted S02 -
S02 conversion to sulfur is calculated by difference as
follows:
S02 Conversion to S = 100 - (R^ + R2 + R3 +
4.4.8 The data in Exhibit No. 4-4 indicate that at a
temperature of 1400°F, 3 seconds contact time with
10% water, and 5 seconds contact time with 0% water, and
at 1500°F and 1 second contact time, greater than 95% CH4
efficiency is attained. About 60-70% S02 is converted to
sulfur and about 10-20% S02 remains unreacted.
4.4.9 A plot of varying temperatures against Cm efficiency
and S02 conversion at fixed contact time of 3 seconds
and 0% H20 in feed is shown in Exhibit No. 4-5. This plot
indicates that at 1450°F about 99% CH4 efficiency and about
90% S02 conversion should be obtainable at the said conditions.
Exhibit No. 4-6 shows a plot of contact time versus ClU
efficiency and S02 conversion at 1500°F and 10% H20 wherein
4-4
-------
Heating Tape
Burners
Thermocouples
for =>
Catalyst
Bed Temp.
STACK
TG&x&see
t '
/ t
\
\ \
Fee
Samp
Pol
p
JJULajL'
1 I
I
d I
tie *
nt C
£
a
J
;
)
) Immersion
Heater
J —
I Ti
j
>
>
;
>
?
^
3
3
h^
--
:
-
^-
1
^
'h
}
~-
—
---
h r
Thermometer
©^
r
The rm-O-Watch
REACTOR
Thermocoup
for 3-Zone
Furnace
Exit
Sample
Point
Chimney
Stopcock
METHANE H20
SATURATOR
Thermowell
THREE-ZONE
FURNACE
Mixing
Chamber
Insulation
Catalyst
~"'ssf7^rr/, 7si /./.j.
-Heating Element
Heating Tape
— Insulation
FLCWMETERS
CONTROL
PANEL
EXHIBIT NO. 4-1
Strong SOg Reduction
Equipment Set-Up
S02
X
-------
about 1 second appears to be the optimum contact time
corresponding to a temperature of 1500°F (97% CH4 efficiency
and 83% S02 conversion).
4.4.10 A comparison between 070 H20 and 10% H20 to determine
the effect of the presence of H20 in the reaction
between strong S02 and CR* is found in Exhibit No. 4-7.
The data indicate no significant variation in CH4 efficiency
and total S02 conversion between feeds containing 0% H20
and 10% H20.
4-5
-------
EXHIBIT NO. 4-3
STRONG SOg/OU REDUCTION
Experimental Data
Run
No.
SSR-1
SSR-2
SSR-3
SSR-4
SSR-5
SSR-6
SSR-3D
SSR-11
SSR-12
SSR-9
SSR-8
SSR-7
Nominal
Temp.
°F
1000
1200
1200
1300
1400
1400
1400
1400
1500
1500
1500
1500
Actual
Contact
Time
sec.
3
3
5
5
3
5
3
5
1/2
1
2
2
H20 in
Feed
%
0
0
0
0
0
0
10
10
10
10
10
0
Gas Composition, Volume 7»
Feed
(Dry Basis)
H2
0
0
0
0
0
0
0
0
0
0
0
OP
0
0
0
0.04
0.05
0
0.07
0.05
0
0.03
0
NP
0.19
0.16
0.15
0.26
0.36
0.12
0.45
0.28
0
0.22
0.19
cm
. __ ur»
33.80
34.07
32.40
33.80
28.80
32.38
40.1
33.65
32.45
32.8
36.6
CO
AN.
0
0
0
0
0
0
0
0
0
0
0
COp
&LYSL
0.10
0.17
0.12
0.01
0
0
0.15
0.11
0
0
0.01
COS
0
0
0
0
0
0
0
0
0
0
0
HpS
0
0
0
0
0
0
0
0
0
0
0
CSp
0
0
0
0
0
0
0
0
0
0
0
SOp
65.5
65.6
67.3
66.0
70.8
67.5
59.3
66.0
67.55
66.9
63.2
Unnormalized Exit
(Dry and S-Free Basis)
HP
0
0
0
0.14
1.59
0.40
1.43
1.15
1.41
1.28
0.93
0.82
OP
0.03
0
0
0.11
0.02
0.06
0.02
0.01
0.02
0.03
0.10
0.01
N2
0.25
0.19
0.18
0.46
0.38
0.51
0.64
0.51
0.30
0.36
0.72
0.29
CH4
34.0
35.0
33.6
30.1
4.88
0.11
1.83
0.64
22.75
1.22
0.02
0.07
CO
0
0.10
0.25
0.77
1.38
0.25
0.63
0.39
1.68
0.60
0.31
0.38
C02
0.43
0.90
2.16
6.1
57.4'
55.89
47.9
59.3
19.4
51.7
59.6
60.1
COS
0
0.02
0.09
0.40
2.83
0.94
1.36
1.14
0.69
1.45
1.22
1.35
. H2S
0
0.09
.0.22
0.68
15.65
11.65
21.7
21.4
5.60
16.08
20.3
23.18
CS2
0
0
0
0.43
6.25
1.55
5.31
1.97
2.52
4.15
0.63
0.83
S02
65.2
63.5
63.6
60.8
19.65
28.64
19.3
13.4
45.6
21.35
16.05
12.96
4-8
-------
EXHIBIT NO. 4-2
STRONG SOg/OU REDUCTION
Theoretical Feed Compositions
(Volume %)
0% HgQ
Component
S02
CH4
Total
Dry Basis
66. 67%
33.33
100.00%
10% H3Q
Component
S02
CH4
H20
Total
Wet Basis
60.0%
30.0
10.0
100.0%
Dry Basis
66.67%
33.33
100.00%
4-7
-------
EXHIBIT NO. 4-4
STRONG SOg/CIU REDUCTION
CH4 and SOg Conversions at: Various
Temperatures, Contact Times and H20 Content in Feed
Run
No.
SSR-1
WbJ&X JL
SSR-2
SSR-3
SSR-4
SSR-5
SSR-6
SSR-10
SSR-11
SSR-12
SSR-9
SSR-8
SSR-7
Nominal
T^1T»T^
lemp.
°F
1000
^ w w W
1200
1200
1300
1400
1400
1400
1400
1500
1500
1500
1500
Contact
T1 1 YYIO
1 JJufc:
Sec.
3
*j
3
5
5
3
5
3
5
1/2
1
2
2
H20 in
*C*Q *"* A
reeo
o
V
0
0
0
0
0
10
10
10
10
10
0
The or.
O/^ / ^^U
\}\J ^ / V^AL4
Ratio
2/1
*•/ *•
2/1
2/1
2/1
2/1
2/1
2/1
2/1
2/1
2/1
2/1
2/1
Ctt*
TJ* -C -C •• /•« n A n iO \T
CiZnciency
2.76
6.75
19.77
92.27
99.54
95.89
98.38
49.99
97.14
99.47
99.41
Actual
SOs/OU
Feed
Ratio
1.82
1.89
1.98
1.90
2.34
1.96
1.97
1.85
2.09
2.06
1.99
To S
2.79
6.45
16.10
63.42
67.81
52.62
68.09
34.54
61.73
69.56
68.58
S02
To COS
SP A/TTrt*
viLALl 1UI
0.03
0.13
0.53
2.04
0.68
1.22
0.91
0.79
1.18
0.96
1.08
Convers:
To H2S
0.14
0.32
0.91
11.31
8.46
19.40
17.12
6.44
13.04
15.91
18.61
Lon
To CS2
0
0
1.15
9.03
2.25
9.50
3.15
5.80
6.73
0.99
1.33
Total
2.96
6.91
18.69
85.80
79.21
82.74
89.28
47.57
82.68
87.42
89.60
Unre-
acted
S02
97.04
93.09
81.31
14.20
20.79
17.26
10.72
52.43
17.32
12.58
10.40
I
vo
-------
CH4 Efficiency
S02 Conversion
EXHIBIT NO. 4-5
CH4 Efficiency and
Conversion vs. Temperature at
3~Sec. Contact Tiro
and 0% HaO in Feed
-------
CIU Efficiency
.
S02 Conversion
EXHIBIT NO. 4-6
CH*. Efficiency and SO
Conversion vs. Contac
Time" atr E500T and
107. EsO in Feed
*
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o
Contact Time, Seconds 4-11
.
-------
EXHIBIT NO. 4-7
STRONG SOg/Cm REDUCTION
Comparison Between 0% H^Q and
10% H20 DaFa
Nominal
Temp.
°F
1400
1400
1400
1400
1500
1500
Contact
Time
Sec.
3
3
5
5
2
2
H20 in
Feed
%
0
10
0
10
0
10
CH4
Efficiency
%
92.27
95.89
99.54
98.38
99.41
99.47
Total S02
Conversion
%
85.80
82.74
79.21
89.28
89.60
87.42
4-12
-------
5. STRONG SOg FROM SMELTER GAS
ASARCO DMA PROCESS
-------
TABLE OF CONTENTS
5.1 INTRODUCTION _ 5-1
5.2 SUMMARY 5-1
5.3 CONCLUSIONS AND RECOMMENDATIONS 5-2
5.4 DMA PROCESS STUDY _ 5-2
5.5 ECONOMIC STUDY 5-11
5.6 REFERENCES 5-13
EXHIBIT 5-1 Flow Sheet, Asarco DMA Process 5-14
EXHIBIT 5-2 DMA Economics Summary 5-15
EXHIBIT 5-3 Equipment Costs - Case 2B _ 5-16
EXHIBIT 5-4 Equipment Costs - Case 2C 5-19
EXHIBIT 5-5 Equipment Costs - Case 2D 5-22
EXHIBIT 5-6 Operating Costs - Case 2fi 5-25
EXHIBIT 5-7 Operating Costs - Case 2C 5-26
EXHIBIT 5-8 Operating Costs - Case 2D 5-2?
EXHIBIT 5-9 Capacity - Cost Relation 5-28
EXHIBIT 5-10 Costs vs. S02 Concentration 5-29
EXHIBIT 5-11 Costs vs Plant Size 5-30
5-i
-------
5.1 INTRODUCTION
5.1.1 Evaluation of direct reduction to sulfur processes for
applicability to S02 pollution abatement in our
Phase I studies resulted in several conclusions. Clearly,
direct reduction was not applicable to power generating stack
gases. In smelter gas studies direct reduction showed
acceptable costs only for roaster type gases which are
relatively strong in S02 and low in oxygen. However,
practically all of the 80s being stacked by smelters is in
the form of a weak gas high in oxygen, typically 1-5% S02
and about 15% oxygen. Direct reduction process costs for
these gases were excessively high. Based on our knowledge
at the time of our Phase I studies, it appeared that only
the low temperature Claus (LTC) process would have
applicability to low S02-high oxygen gases. Early in
our Phase II laboratory studies on the LTC process, there
were indications that assumptions made in our Phase I
evaluations were not valid. This pointed out the need for
an alternate method of handling smelter SO2 emissions.
5.1.2 The low cost of reducing very strong S02 that resulted
from some upstream gathering process, using methane as
the reductant, suggested that the combined cost of concentrating
the SO2 and subsequent reduction to sulfur could provide a
tenable solution to the smelter pollution problem. On that
basis, a survey was made of concentrating schemes that might
be employed upstream of the reduction process. The most
practical of these appeared to be the dimethylaniline (DMA)
sorption process developed by Asarco and used by them in a
20 ton per day S02 plant in Selby, California. Accordingly,
the study reported herein is based on use of the Asarco DMA
process to scrub smelter exits down to 0.045% S02, and deliver
a 97% S02 gas to the reduction plant.
5.2 SUMMARY
5.2.1 Three cases were detailed, based on gases containing
400 NTD of S02 (200 NTD of sulfur), with initial
concentrations of 2.9, 4.5 and 8.0 percent S02.
5.2.2 Factorial extrapolations were made based on costs
developed for the 200 NTD sulfur plants to estimate
the costs for a 100 and a 400 NTD sulfur plant for each of
the gas cases.
5.2.3 These costs were combined with those developed for
methane reduction of a 90% S02-10% H20 gas in our
Phase I studies.
5-1
-------
5.3 CONCLUSIONS AND RECOMMENDATIONS
5.3.1 A combined system of DMA concentration and high
temperature methane reduction of S02 in smelter gas
is highly promising. At a production level of 200 NTD of
sulfur, using the 4.5% S02 feed case, fixed capital totals
out at $6,400,000 and operating costs at $33.51 per short
ton of sulfur. These costs become lower as plant size
increases or feed gas is higher in S02 content.
5.3.2 For the 2.9 to 8.0% S02 range, the combined DMA
concentration-high temperature ClU reduction process
is clearly better than direct CiU reduction. Both fixed
capital and operating costs are lower. This will hold
for the 100 to 400 NTD sulfur capacity levels.
5.3.3 The DMA plant is costly due to use of SS 316
throughout as a proven material. There are probably
opportunities to reduce capital cost by substitution of
reinforced plastics, liners, and the like. The savings in
fixed charges could be significant.
5.3.4 In these studies, a steady flow and constant composition
to the DMA process were assumed. In the real smelter
situation, some modification in the design and operation of
the DMA unit will be needed to accomodate the fluctuating
gas volume and SOa content.
5.3.5 Work should be done on developing and optimizing a
reduction process for 100% S02 that will represent
improvements in operability and economics over heretofore
visualized processes.
5.4 DMA PROCESS STUDY
5.4.1 General
5.4.1.1 Because of the high cost of sulfur recovery
by reduction of weak smelter gases (2.9% to
4.5% S02), a survey was made of concentrating schemes that
might be employed upstream of the reduction process. The
most practical of these appears to be the dimethylaniline
(DMA) sorption process used by Asarco at their plant in
Selby, California (References 1 and 2). This plant and its
performance has been reported in sufficient detail to allow
a scaleup from its 10 tons per day sulfur equilvalent from
a 5% S02 feed gas to the 100 to 400 tons/day range reported
here.
5.4.1.2 Three cases have been worked up, two based
on the 2B and 2c gases defined in the
PH 22-68-24 Phase I report and a third gas designated 2o.
The gas compositions and the volume equivalent to 400 tons
SOs/day are summarized as follows:
5-2
-------
Table 5-1
Gas Designation
Volume 7, 80s
Volume % 02
Volume % N2 and Inert
SCFM
2D
8.0
8.8
83.2
39,400
2B
4.5
16.5
79.0
70,000
2C
2.9
14.3
82.8
108,600
5.4.1.3 The Selby unit employed lead equipment
throughout. More recent practice utilizes
stainless steel type 316, therefore stainless is specified
herein for lower costs. Exhibit 5-1 is a flow diagram of
the process including equipment for cooling and cleaning
the input smelter gas.
5.4.1.4 The sizing of the units and the estimation of
cooling and heating duties are based on the
following data.
Table 5-2
SQg Solubility in DMA
Gas
2D
2B
2C
Equilibrium (68°F)
340 gpl
225 gpl
150 gpl
Operating
204 gpl
135 gpl
90 gpl
Table 5-3
SP. Heat of DMA
32-68°F
32 °F
68 °F
77°F
0.416
0.403
0.430
0.440
Heat of solution of S02 in DMA at 72°F - 22100 BTU/# mole
S02. Sp. gravity of DMA is 0.956 at 68°F.
5-3
-------
Table 5-4
Viscosity of DMA
°F
50
68
86
104
122
Centipoise
1.69
1.40
1.17
1.04
0.91
Both the density and viscosity of DMA are very close to water,
so water ratings are used for DMA pump duty.
5.4.2 Gas Cooler and Cleaner
5.4.2.1 This is a packed tower designed according
to Chemical Engineers' Handbook, Perry,
Ed. Ill, pages 680-681. The gas is taken at 500°F and
carrying 0.2 grains dust/scf. If the smelter gas is
substantially hotter than 500°F we assume it would be
dropped to 500°F in a waste heater boiler. The dust loading
of 0.2 grains/scf will give a solids content of 0.046% in
the water leaving the unit. Water enters the tower at 65°F
and discharges at 180°F. The solids are settled out in a
pond. The smelter gas is cooled to 68°F.
Table 5-5
Gas Cooler Design Conditions
Cooling water to tower
Cooling water from tower
Tower diameter - feet
Gas —
gpm
gpm
— > 2D
303
296
15.0
2B
530
517
20.0
2C
823
804
25.0
5.4.2.2 An arbitrary packed depth of 50 feet is used;
this may be more than necessary. A packing
of 2" dumped rings will have a pressure drop of 0.40" H20/ft
or 20" H20 on a 50 foot bed. Indicated operating conditions
will be safely below the flood point.
5-4
-------
5.4.3 DMA Absorber - Scrubber
5.4.3.1 This bubble cap absorber is sized according
to Perry, Ed. Ill, page 579 and conditions
are as follows:
Liquid Density - 0.98 sp. gr.
Tray Spacing - 24"
Liquid Seal - 1.5"
Superficial Velocity - 4.55 ft/sec
Absorber
Soda Scrubber
Acid Scrubber
TOTAL
8 trays
2 trays
9 trays
19 trays
Assume pressure drop of 2.0" water/tray, then total pressure
drop - 38"water.
Table 5-6
Absorber Design Conditions
Cln.fi ..— -..S
Lbs. pregnant DMA/min.
Lbs. SOa/min.
Lbs. DMA/min.
Gals. DMA/min.
Absorber Dia. - ft
Absorber Height - ft
2D
2881.5
561.9
2319.6
292
15.3
50.0
2B
4159.2
561.5
3597.7
452
19.5
50.0
2C
6238.8
561.5
5677.3
713
23.9
50.0
5.4.4 Absorber Intercoolers
5.4.4.1 The intercoolers are numbered on Exhibit 5-1,
' No detailed information on the tray tempera-
tures in the Selby unit was given. In this design a larger
fraction of the S02 absorption is taken on the lower plates
where the partial pressure of S02 is higher. This reduces
the heat release on the upper trays where it is important
to have a low DMA temperature on trays scrubbing the
leanest gas.
Cooling Water - 65°F
U - 100 BTU/hr - sq. ft. - QF
5-5
-------
5.4.4.2 Table 5-7 following carries a summary of
intercooler parameters.
5.4.5 Intercooler Recirculation Pumps
5.4.5.1 Pumps for the Indicated volume of DMA are
rated as if DMA were water. Required head
is equal to the pressure loss in the exchangers and lines.
5.4.6 Dilute HgSCU and Dilute Soda Supply
5.4.6.1 The reagent requirements and the supply
tank size for a three-day supply of each
dilute reagent is an follows. DMA make may be from a small
storage tank or from drums.
Table 5-8
Reagent
Soda Ash #/day
H2S04 (100%) #/day
DMA #/day
3 day supply 20% reagent, gal.
2D
8000
9000
250
11000
2B
14200
16000
440
20000
2C
22200
25000
590
30000
5.4.7 Smelter Gas Blower
5.4.7.1 Pressure drop of gas cooler is 20" H20,
plus the absorber drop of 38" H20 gives
a total to 58". Horsepower at 60% efficiency is tabulated
below.
Table 5-9
cfm (68°F)
Horsepower
2D
42300
640
2B
77500
1170
2C
120000
1820
5.4.8 Phase Separator and Surge Tanks
5.4.8.1 Based on 15 minutes retention in the phase
separation system, the following tank sizes
are required.
Table 5-10
DMA Surge tank and separator - gal
Stripper water surge tank - gal
2D
5200
1000
2B
7500
1100
2C
11000
1200
5-6
-------
Table 5-7
Intercooler Parameters
Cooler No.
Gas Type
DMA #/min
DMA in *F
DMA out *F
Water GPM
Water in *F
Water out °F
Log mean AT, °F
BTU/hr x 10 6
Exchanger
Area ft
1
2D
2880
81
68
198
65
75
4.35
0.99
2280
2B
4160
77
68
394
65
70
4.75
0.99
2090
2C
6240
73
68
547
65
68
3.95
0.82
2080
2
2D
2880
116
68
730
65
75
12.3
3.66
2980
2B
4160
102
68
745
65
75
10.9
3.74
3420
2C
6240
91
68
755
65
75
7.82
3.78
4720
3
2D
2880
106
68
580
65
75
11.8
2.89
2450
2?
4160
94
68
570
65
75
8.7
2.85
3270
2C
6240
84
68
528
65
75
5.58
2.64
4720
4
2D
28 8C
98
68
465
65
75
10.3
2.28
2220
2B
4160
87
68
417
65
75
6.5
2.09
3200
2C
6240
85
68
570
65
75
5.82
2.80
4820
5
2D
2880
80
68
182
65
75
3.94
0.92
2320
2B
4160
77
68
394
65
70
4.75
0.99
2090
2C
6240
73
68
547
65
68
3.95
0.82
2080
Oi
I
•vj
-------
5.4.9 Stripper-Regenerator-Rectifier
5.4.9.1 This unit is sized for a S02 rate of 400 tons/
day and is the same for all cases. The unit
is based on a 1:1 ratio of steam to SQ2 leaving the rectifier
section. It is assumed that the ratio of steam to S02 is
2:1 in the stripping section. The size (dia.) is based on
the stripper duty and is somewhat oversize for the regenerator
and rectifier sections.
Regenerator - 7 trays
Stripper - 14 trays
Rectifier - 5 trays
TOTAL 26 trays
Design basis for stripper section:
Vapor Volume - 9750 scfm
Vapor Temperature - 210°F
Liquid density at 210 *F - 0.95 sp. gr.
Plate Spacing - 18"
Liquid seal - 1.5"
Superficial Velocity - 3.44 ft/sec.
Then:
Column diameter - 9.1 ft.
Overall Column Height - 45 ft.
5.4.10 SQg Product Gas Scrubber - Unit 11
5.4.10.1 The unit serves to remove any traces of DMA
from the final S02 product. It may not be
necessary on gas being reduced to sulfur but it is necessary
where the gas is being converted to pure dry S02 for sale
as such. This scrubber is included assuming the recovery of
DMA justifies its cost. The unit is 6.0 ft. dia. and may be
about 10 ft. packed depth. With 10' of packing, the pressure
drop will be about 3.1" H20 using 1 1/2" dumped rings.
Sizing is the same for all three cases.
5.4.11 Stripper Condenser - Unit 7
5.4.11.1 This unit is the same for all three cases.
The hot S02-H20 vapor leaving the S02
rectifier is cooled. The water condensate formed is returned
to the top plate of the rectifier.
5-8
-------
Table 5-11
Stripper Condenser Design - Unit 7
Gas in
Gas and Water out
Cooling H20 in
Cooling H20 out
Log mean AT
BTU/hr
Transfer Surface
Cooling Water
- 179CF
- 85 °F
- 65 °F
- 125 °F
- 26.2°F ,
- 6.1 x 106
- 4680 ft2
- 205 gpm
Table 5-12
DMA Exchanger - Unit 8
Pregnant DMA in #/min
Net DMA in #/min
Pregnant DMA in °F
Pregnant DMA out QF
DMA in °F
DMA out °F
BTU/hr exchange x 1C6
Log mean AT °F
Exchanger area ft2, basis U " 100
BTU/hr ft2 °F
2D
2880
2320
68
178
210
100
7.55
32.0
2360
2B
4160
3600
68
178
210
100
11.3
32.0
3700
2C
6240
5680
68
178
210
100
17.3
32.0
5420
5-9
-------
Table 5-13
DMA Cooler - Unit 6
Cooling water in °F
Cooling water out °F
Log Mean AT °F
DMA #/min
Water gpm
BTU/hr x 10 6
Exchanger area ft2, basis U ° 100
DMA in °F
DMA out °F
2D
65
75
10.5
2320
390
1.95
1850
100
68
2B
65
75
10.5
3600
607
3.02
2880
100
68
2C
65
75
10.5
5680
955
4.77
4550
100
68
5.4.12 Regenerator Reboiler - Unit 10
5.4.13 Stripper Reboiler - Unit 9
5.4.13.1 These units are thermosyphon reboilers
operating on a water phase. The stripper
reboiler water phase probably will carry a small amount of
entrained DMA. Conditions taken are as follows:
Steam side - 100 psig steam at 338°F
Water phase - 220°F
U - 200
AT - 118°F
Gives 42.3 ft2 for 106BTU/hr
Table 5-14
BTU/hr x 10 6
Steam #/min
Transfer area ft2
2D
Regen.
11.9
225
500
Strip
13.2
250
560
2B
Regen.
12.2
231
520
Strip
14.4
273
610
2C
Regen.
12.6
240
530
Strip
16.1
304
680 -
5-10
-------
Steam Demand - 100 psig
Table 5-15
Steam - Ibs/hr.
# s team/ # S02 produced
2v
28500
0.85
2B
30240
0.90
1C
32600
0.97
5.4.14 65°F Cooling and Process Water
Table 5-16
Gas Cooler and Cleaner gpm
Intercoolers gpm
DMA Cooler gpm
Rectifier Condenser gpm
Total gpm
2D
296
2155
390
205
3046
2B
530
2520
607
205
4612
2C
823
2950
955
205
4933
5.4.15 SO? Yield %
2D
2B
2C
99.4
99.0
98.4
Gas exit - 0.045% S02
5.5 ECONOMIC STUDY
5.5.1 Based on the foregoing process design, plants to
recover 400 NTD 97% S02 gas from feed gases ranging from
2.9% to 8.0% S02 according to the flow sheet Exhibit 5-1
were estimated. The results are summarized in Exhibit 5-2,
supported by details in Exhibits 5-3 through 5-8 following.
5.5.2 Factorial extrapolations were made from the 400 NTD
SQz level (200 NTD equivalent sulfur) to cover the
range 100 to 400 NTD sulfur capacity. Fixed capital and
operation costs are summarized in Exhibit 5-9 and the complete
5-11
-------
economic picture? portrayed in the plots marked Exhibit 5-10
(where vol. '/„ S0;. is the parameter) and Exhibit 5-Li where
NTD sulfur equivalent is the parameter.
5.5.3 Besides developing S02 concentration costs via the
DMA process, a principal purpose of this study was
to compare the cost of direct methane reduction of "as is"
weak smelter gases with the combined cost of DMA concenLration
followed by methane reduction.
5.5.4 It was earlier established in our phase I studies
that a 90% 80s stream could be reduced with methane
for about $11/NT sulfur (200 NTD sulfur capacity), whereas
operation on weak gases incurred costs ranging from $34 to
$47 per NT sulfur. Tables 5-17, 5-18 and 5-19 were prepared
to permit this comparison.
Table 5-17
Basis 2QQ NTD S Equivalent in Feed
Gas No.
2D
2B
2C
SO;;
8 . 0%
4.5%
2.9%
DMA Process
FC. S MM
$3.9
5.0
6.0
OC/NT
$17.02
$22.57
$28.02
Asarco CH4 Red n*
FC, $ MM
$1.40
$1.40
$1.40
OC/NT
$10.94
$10.94
$10.94
Total
PC, $ MM
$5.3
6.4
7.4
OC/NT
$27.96
$33.51
$38.96
All basis 90% S02, 10% H20 Feed.
Table 5-18
Basis 400 NTD S Equivalent in Feed
Gas No.
2D
2B
2C
SOp.
8.0%
4.5%
2.9%
DMA Process
FC, $ MM
$5.9
7.6
9.0
OC/NT
$12.12
$16.39
$20.56
Asarco CH^ Red'n
FC, $ MM
$2.10
$2.10
$2.10
OC/NT
$ 8.59
$ 8.59
$ 8.59
Total
FC, $ MM
8.0
9.7
11.1
OC/NT
$20.71
$24.98
$29.15
Table 5-19
Other Methods
Gas No.
2B
2D
S02
4.5%
8.0%
Process
Asarco Type CH4 Red.
No Data
FC, $ MM
$9.15
-
OC/NT
$46.83
-
NTD S in Feed
200
-
5-12
-------
5.6 REFERENCES
1. Kohl and Riesenfeld, "Gas Purification", pages 204
through 209.
2. Ind. Eng. Chem. 42, No. 11, pages 2253-2258, "Liquid
S02 from Waste Smelter Gasesfl, E. P. Fleming & T. C. Fitt
5-13
-------
•feter. 65 •?
C*a
Cooler
and
Cleaner
aelter Ca«
500 *F
20t
L
To
Atnoaphere
Dilute Acid
Supply
201
!UaCO
,1
Dilute Soda
Supply
Soda
Scrubber
Absorber
Booster Fan
180"F Water
to Pond
EXHIBIT 5-1
STtOHC SOa FMi SMELTgR GAS
400 TOM SOa per Day
II"
— F=-:
Abiorber
Heater
Stripper
Water to
Waste 210°F|
5-14
-------
EXHIBIT 5-2
DMA Economics Summary
"Order of Magnitude11 estimates of capital cost to produce
400 NTD S02 gas (200 NTD sulfur equivalent) at 97.1 vol.% purity
were prepared] supporting details are given in Exhibits 5-3,
5-4 and 5-5 following. The results are summarized below.
Table 5-20
Case
2D
2B
2C
Smelter Gas Feed Composition
7o SO*
8.0
4.5
2.9
% 02
8.8
16.5
14.3
% Np + Inert s
83.2
79.0
82.8
Purchased
Equipment
$ 900,000
$1,352,000
$1,825,000
Battery Limits*
Fixed Capital
$ 3,900,000
$ 5,000,000
$ 6,000,000
* The Battery Limits capital excludes nickel surcharge on
all stainless steel equipment, includes approximately
25% contingency. No buildings, site preparation, service
facilities or effluent control facilities have been
included. It is assumed that all utilities at appropriate
conditions are available at the battery limits. Piping,
instrumentation, electrical gear and overheads were
factored into the totals basis purchased equipment cost.
Utility requirements are tabulated below.
Table 5-21
Steam, M Ibs
Process & Cooling Water, M gals
Electricity, KWH
Feed Gas, SCFM at 400 NTD S02
Per NT SOg Product
Case 2D
1.70
11.0
33.5
39,400
Case 2B
1.80
16.7
59.1
70,000
Case 2c
1.94
17.8
92.9
108,600
5-15
-------
EXHIBIT 5-3
Recovery of 97% SOg Vapor From
4.5% SOg Bearing Smelter Gases
DMA_Process - Case gB
Summation of Major Equipment
1. Gas Cooler and Cleaner - Tower of 316 stainless steel, scrubbing
and cooling smelter gas from 500°F to 68°F. 20 ft. diameter
with 50 ft. depth of ceramic packing $340,000
2. DMA Absorber Scrubber - Bubble cap tower of 316 stainless
steel 19 ft.-6 in. diameter by 50 ft. height with 8 absorbing
trays, 2 soda scrubbing trays and 9 acid
scrubbing trays $430,000
3. Absorber Intercoolers - Each a 316 stainless steel exchanger.
#1 2090 sq. ft. $ 20,200
#2 3420 sq. ft. $ 33,300
#3 3270 sq. ft. $ 31,000
#4 3200 sq. ft. $ 31,000
#5 2090 sq. ft. $ 20,200
4. Recirculating Pumps - Five centrifugal pumps, each of
316 stainless steel, delivering 520 gpm at 40 ft. TDH with 10
HP TEFC motor. Total for five (5) $ 8,000
5. Dilute Acid Supply Tank - a 20,000 gal. tank of 316 stainless
steel $ 22,000
6. Dilute Soda Supply Tank - a 20,000 gal. tank of 316 stainless
steel $ 22,000
7. DMA Supply Tank - a 200 gal. tank of carbon steel - $ 400
8. Smelter Gas Blower - To deliver 77,500 cfm at 68°F against
58 in. water, made of 316 stainless steel s 46,400
Drip-proof motor, 1200 HP $ 18,400
9. Phase Separator - A 7,500 gal. tank of 316 stainless
steel $ 11,900
10. DMA Surge Tank - A 7,500 gal. tank of 316 stainless
steel $ 11,900
11. Separator Tank - A 7,500 gal. tank of 316 stainless
steel $ 11,900
12. Stripper Water Surge Tank - A 1,100 gal. tank of
316 stainless steel - $ 2,100
5-16
-------
13. Stripper Regenerator Rectifier - Bubble cap tower of 316
stainless steel 9 ft. diameter by 45 ft. height with
7 regenerating trays, 14 stripping trays, and
5 rectifying trays - $ 135,000
14. SOa Product Gas Scrubber * Tower of 316 stainless steel
6 ft. diameter with 10 ft. of ceramic packing $ 23,000
15. Stripper Condenser - A 316 stainless steel exchanger
with 4680 sq. ft. transfer area $ 45,200
16. DMA Exchanger - A 316 stainless steel exchanger with
3700 sq. ft. transfer area $ 36,000
17. DMA Cooler - A 316 stainless steel exchanger with
2880 sq. ft. transfer area - -- $ 27,400
18. Regenerator Reboiler - A 316 stainless steel unit
with 520 sq. ft. transfer area and up to 100 psig
steam on the shell side — --- $ 6,600
19. Stripper Reboiler - A 316 stainless steel unit with
610 sq. ft. transfer area $ 7,400
20. Collecting Tank - A 200 gal. tank of 316 stainless
steel $ 1,200
21. Gas Cooler Pump - A 316 stainless steel centrifugal
pump to deliver 700 gpm at 100 ft. TDH with 30 HP
TEFC motor $ 2,200
22. 20% Acid Pump - A 316 stainless steel centrifugal
pump to deliver 20 gpm at 30 ft. TDH with a 2 HP
TEFC motor $ 700
23. 20% Soda Pump - 316 stainless steel centrifugal pump
to deliver 20 gpm at 30 ft. TDH with a 2 HP TEFC
motor $ 700
24. Stripper Feed Pump - A 316 stainless steel centrifugal
pump to deliver 500 gpm at 100 ft. TDH with a 30 HP
TEFC motor $ 1,900
25. DMA Makeup Pump - A 316 stainless steel centrifugal
pump to deliver 20 gpm at 30 ft. TDH with a 2 HP
TEFC motor $ 700
26. Separator feed Pump - A 316 stainless steel centrifugal
pump to deliver 20 gpm at 30 ft. TDH with a 2 HP
TEFC motor - $ 700
5-17
-------
27. Recycle Pump - A 316 stainless steel centrifugal pump
to deliver 500 gpm at 100 ft. TDH with a 30 HP
TEFC motor $ 1,900
28. Stripper Bottoms Pump - A 316 stainless steel
centrifugal pump to deliver 20 gpm at 30 ft. TDH with
a 2 HP TEFC motor - -- $ 700
TOTAL PURCHASED
EQUIPMENT
5-18
-------
EXHIBIT 5-4
Recovery of 97% SQg Vapor From
2.9% SOg Bearing Smelter Gases
DMA~Process -f Case ZC
Summation of Major Equipment
1. Gas Cooler and Cleaner - Tower of 316 stainless steel,
scrubbing and cooling smelter gas from 500°F to 68°F
25 ft. diameter with 50 ft. depth of ceramic
packing $ 478,000
2. DMA Absorber-Scrubber - Bubble cap tower of 316 stainless
steel 24 ft. 6 in. diameter by 50 ft. height with
8 absorbing trays, 2 soda scrubbing trays and 9 acid
scrubbing trays $ 614,000
3. Absorber Intercoolers - Each a 316 stainless steel
exchanger.
#1 2080 sq. ft. $ 20,100
#2 4720 sq. ft. $ 45,600
#3 4720 sq. ft. $ 45,600
#4 4820 sq. ft. $ 46,500
#5 2080 sq. ft. $ 20,100
4. Recirculating Pumps - Five centrifugal pumps, each of 316
stainless steel, delivering 800 gpm at 40 ft. TDH with
15 HP TEFC motor. Total for five (5) $ 10,800
5. Dilute Acid Supply Tank - A 30,000 gal. tank of 316
stainless steel $ 26,400
6. Dilute Soda Supply Tank - A 30,000 gal. tank of 316
stainless steel $ 26,400
7. DMA Supply Tank - A 200 gal. tank of carbon steel - $ 400
8. Smelter Gas Blower - To deliver 120,000 cfm at 68°F
against 58 in. water, made of 316 stainless steel- $ 95,100
Drip-proof Motor, 1800 HP $ 25,300
9. Phase Separator - An 11,000 gal. tank of 316 stainless
steel $ 14,800
10. DMA Surge Tank - An 11,000 gal. tank of 316 stainless
steel -- - - $ 14,800
11. Separator Tank - An 11,000 gal. tank of 316 stainless
steel $ 14,800
12. Stripper Water Surge Tank - A 1,200 gal. tank of 316
stainless steel - $ 2,300
5-19
-------
13. Stripper Regenerator Rectifier - Bubble cap tower of 316
stainless steel 9 ft. diameter by 45 ft. height with
7 regenerating trays, 14 stripping trays, and
5 rectifying trays $ 135,000
14« SO^ Product Gas Scrubber - Tower of 316 stainless steel
6 ft. diameter with 10 ft, of ceramic packing $ 23,000
15. Stripper Condenser - A 316 stainless steel exchanger with
4680 sq. ft. transfer area $ 45,000
16. DMA Exchanger - A 316 stainless steel exchanger with
5420 sq. ft. transfer area $ 50,000
17. DMA Cooler - A 316 stainless steel exchanger with 4550
sq. ft. transfer area $ 43,900
18. Regenerator Reboiler - A 316 stainless steel unit with
530 sq. ft. transfer area and up to 100 psig steam on
the shell side - $ 6,700
19. Stripper Reboiler - A 316 stainless steel unit with
680 sq. ft. transfer area $ 8,200
20. Collecting Tank - A 200 gal. tank of 316 stainless
steel - $ 1,200
2l. Gas Cooler Pump - A 316 stainless steel centrifugal
pump to deliver 1100 gpm at 100 ft. TDH with 50 HP
TEFC motor $ 2,800
22. 2070 Acid Pump - A 316 stainless steel centrifugal pump
to deliver 20 gpm at 30 ft. TDH with a 2 HP
TEFC motor - $ 700
23. 20% Soda Pump - A 316 stainless steel centrifugal pump
to deliver 20 gpm at 30 ft. TDH with a 2 HP TEFC
motor $ 700
24. Stripper Feed Pump - A 316 stainless steel centrifugal
pump to deliver 750 gpm at 100 ft. TDH with a 40 HP
TEFC motor $ 2,500
25. DMA Makeup Pump - A 316 stainless steel centrifugal
pump to deliver 20 gpm at 30 ft. TDH with a 2 HP TEFC
motor $ 700
26. Separator Feed Pump - A 316 stainless steel centrifugal
iump to deliver 20 gpm at 30 ft. TDH with a
HP TEFC motor $ 700
5-20
-------
27. Recycle Pump - A 316 stainless steel centrifugal pump to
deliver 750 gpra at 100 ft. TDK with a 40 HP TEFC
motor $ 2,500
28. Stripper Bottoms Pump - A 316 stainless steel centrifugal
pump to deliver 20 gpm at 30 ft. TDH with a 2 HP TEFC
motor $ 700
TOTAL PURCHASED EQUIPMENT $1,825,500
5-21
-------
EXHIBIT 5-5
Recovery of 97% SO- Vapor From
8". 07, SOL Bearing Smelter Gasc-s
DMA~Process - Case 2"D
Summation of Major Equipment
1. Gas Cooler and Cleaner - Tower of 316 stainless steel,
scrubbing and cooling smelter gas from 500°F to 6801>'.
15 ft. diameter with 50 ft. depth of ceramic packing
$ i'-n,ooo
2. DMA Absorber-Scrubber - Bubble cap tower of 316 stainless
steel 15 ft. 4 in. diameter by 50 ft. height with 8
absorbing trays, 2 soda scrubbing trays and
9 acid scrubbing trays $ 201,000
3. Absorber Intercoolers - Each a 316 stainless steel cxchnnj.;cr
#1 2280 sq. ft. $ 22.10'
#2 2980 sq. ft. $ .'.S,30(;
#3 2450 sq. ft. $ 23, JOG
#4 2200 sq. ft. $ 21,30')
#5 2320 sq. ft. $ 112,40'j
4. Recirculating Pumps - Five centrifugal pumps, each of
316 stainless steel, delivering 360 gpm at 40 ft.
TDH with 7.5 HP TEFC motor $ 8,000
5. Dilute Acid Supply Tank - A 11,000 gal. tank of
316 stainless steel $ 14,800
6. Dilute Soda Supply Tank - A 11,000 gal. tank of
316 stainless steel $ 14,800
7. DMA Supply Tank - A 200 gal. tank of carbon steel- $ 400
8. Smelter Gas Blower - To deliver 42,300 cfm at 68°F
against 58 in. water, made of 316 stainless steel- $ 42,800
Drip-proof motor, 600 HP $ 9,300
9. Phase Separator - A 5,200 gal. tank of 316 stainless
steel $ 9. 700
10. DMA Surge Tank - A 5,200 gal. tank of 316 stainless
steel $ 9.700
11. Separator Tank - A 5,200 gal. tank of 316 stainless
steel - $ 9,700
5-22
-------
12. Stripper Water Surge Tank - A 1,000 gal. tank of 316
stainless steel $ 2,100
13. Stripper Regenerator Rectifier - Bubble cap tower of
316 stainless steel, 9 ft. diameter by 45 ft. height with
7 regenerating trays, 14 stripping trays, and
5 rectifying trays $ 135,000
14. S02 Product Gas Scrubber - Tower of 316 stainless steel
6 ft. diameter with 10 ft. of ceramic packing $ 23,000
15. Stripper Condenser - A 316 stainless steel exchanger with
4680 sq. ft. transfer area $ 45,200
16. DMA Exchanger - A 316 stainless steel exchanger with
2360 sqo ft. transfer area $ 22,800
17. DMA Cooler - A 316 stainless steel exchanger with
1850 sq. ft. transfer area $ 17,900
18. Regenerator Reboiler - A 316 stainless steel unit with
500 sq. ft. transfer area and up to 100 psig steam
on the shell side $ 6,300
19. Stripper Reboiler - A 316 stainless steel unit with
560 sq. ft. transfer area $ 7,050
20. Collecting Tank - A 200 gal. tank of 316 stainless
steel $ 1,200
21. Gas Cooler Pump - A 316 stainless steel centrifugal
pump to deliver 400 gptn at 100 ft. TDH with 20 HP
TEFC motor $ 1,750
22. 20% Acid Pump - A 316 stainless steel centrifugal pump
to deliver 20 gpm at 30 ft. TDH with a 2 HP TEFC
motor $ 700
23. 20% Soda Pump - A 316 stainless steel centrifugal pump
to deliver 20 gpm at 30 ft. TDH with a 2 HP TEFC
motor $ 700
24. Stripper Feed Pump - A 316 stainless steel centrifugal
pump to deliver 350 gpm at 100 ft. TDH with a 20 HP
TEFC motor - $ 1,350
25. DMA Makeup Pump - A 316 stainless steel centrifugal
pump to deliver 20 gpm at 30 ft. TDH with a 2 HP
TEFC motor $ 700
26. Separator Feed Pump - A 316 stainless steel centrifugal
pump to deliver 20 gpm at 30 ft. TDH with a 2 HP
TEFC motor - — $ 700
5-23
-------
27. Recycle Pump - A 31ft stainless steel centrifugal pump
t<. deliver 350 j.-pni at LOO ft. TDH wiLh a 20 HI'
TKFC moLor S \ ,.}')()
28. Stripper Bottoms I'uinp - A 316 stainless steel centrifugal
pump to deliver 20 gpm at 30 ft. TDIl with a 2 up TEFC
motor $ 700
TOTAL PURCHASED EQUIPMENT $ 1JOO,000
i-24
-------
EXHIBIT
5-6
OPERATING COSTS
PROCESS:
DMA Sorption
CASE NO.
2B
Fixed Capital, Battery Limits Basis
Sulfur Production NTD x 330 Days
$5,000,000
66,000 NTY
Line
No.
1.
2.
3.
4.
5.
6. a
b
c
d
e
7.
8.
9.
0.
1.
2.
3.
4.
5.
6.
7.
8.
9.
0.
1.
2.
3.
Cost Items
DMA
Soda Ash
HSS04, 100%
Total Raw Materials
Direct Labor
Supervision
Maintenance
Plant Supplies
Electricity
Cooling Water
Boiler Feed Water
Fuel Gas
Steam
Total Utilities
Other Direct (Catalyst)
TOTAL DIRECT
Payroll Burden
Plant Overhead
Pack & Ship
Waste Disposal
Other Indirect
TOTAL INDIRECT
Depreciation
Taxes
Insurance
Other Fixed
TOTAL FIXED
TOTAL OPERATING
Credits
NET OPERATING
Used
Per NT
2.2 Ibs
0.036 NT
0.04 NT
-
8750 MR
-
5% FC
10?,,(L2+L3)
118 KWH
33.4 MG
MG
MCF
3.6 M Ib
-
-
30% L2
70%(L2+L3)
1000 MH
-
-
10% FC
2% FC
1% FC
-
-
-
-
Cost
Per Unit
0.20
$40.
$30.
-
$3.50
$10,000
-
-
$.01
.05
.25
.30
.50
-
-
-
-
3.90
-
-
-
-
-
-
-
-
-
Per NT
Sulfur
$ .44
1.44
1.20
$3.08
.47
.08
3.79
.05
1.18
1.68
—
1.80
4.66
-
$12.13
.14
'.38
-
.06
-
$ .58
7.58
1.52
.76
-
$ 9.86
$22.57
-
$22.57
Per Year
$M
$ 29.0
95.0
79.0
$ 203.0
31.0
5.0
250.0
3.6
78.0
111.0
_
118.8
307.8
-
$ 800.5
9.3
25.2
-
3.9
-
$ 38.4
500.0
100.0
50.0
-
$ 650.0
$1,488.9
-
$1,489.
5-25
-------
EXHIBIT
5-7
.OPERATING COSTS
PROCESS:
DMA Sorption
CASE NO.
2C
Fixed Capital, Battery Limits Basis
Sulfur Production NTD x 330 Days
$6,000,000
66,OOONTY
Line
No.
1.
2.
3.
4.
5.
6. a
b
c
d
e
7.
8.
9.
0.
1.
2.
3.
4.
5.
6.
7.
8.
9.
0.
1.
2.
3.
Cost Items
DMA
Soda Ash
Hr_.S04 100%
Total Raw Materials
Direct Labor
Supervision
Maintenance
Plant Supplies
Electricity
Cooling Water
Boiler Feed Water
Fuel Gas
Steam
Total Utilities
Other Direct (Catalyst)
TOTAL DIRECT
Payroll Burden
Plant Overhead
Pack & Ship
Waste Disposal
Other Indirect
TOTAL INDIRECT
Depreciation
Taxes
Insurance
Other Fixed
TOTAL FIXED
TOTAL OPERATING
Credits
NET OPERATING
Used
Per NT
2.95 Ibs
0.06 NT
0.063 NT
8750 MH
-
5% FC
10%(L2+L3)
186 KWH
35.6 MG
MG
MCF
3.88 M lb
30% L2
707o(L2+L3)
1000 MH
10% FC
2% FC
1% FC
-
Cost
Per Unit
$ .20
$ 40.
$ 30.
$3.50
$10,000
-
-
$.01
.05
.25
.30
.50
;
-
3.90
- •
-
-
-
Per NT
Sulfur
$ .59
2.40
1.89
$ 4.88
.47
.08
4.55
.05
1.86
1.78
1.94
5.58
-
$15.61
.14
'.38
-
.06
$ .58
9.10
1.82
.91
-
$11.83
$28.02
$28.02
Per Year
$M
$ 39.0
158.5
125.0
$ 322.5
31.0
5.0
300.0
3.6
122.5
117.0
128.0
367.5
-
$ 1.029.6
9.3
25.2
-
3.9
$ 'M . A
600 . 0
120.0
60.0
-
$ 780.0
$ 1,848.0
-
$ 1,848.
-------
EXHIBIT
5-8
OPERATING COSTS
PROCESS:
DMA Sorptiim
CASE NO.
2D
Fixed Capital, Battery Limits Basis
Sulfur Production NTD x 330 Days
$3,900,000
Line
No.
1.
2.
3.
4.
5.
6. a
b
c
d
e
7.
8.
9.
0.
1.
2.
3.
4.
5.
6.
7.
8.
9.
0.
1.
2.
3.
Cost Items
DMA
Soda Ash
H2S04, 100%
Total Raw Materials
Direct Labor
Supervision
Maintenance
Plant Supplies
Electricity
Cooling Water
Boiler Feed Water
Fuel Gas
Steam
Total Utilities
Other Direct (Catalyst)
TOTAL DIRECT
Payroll Burden
Plant Overhead
Pack ft. Ship
Waste Disposal V
Other Indirect i'
TOTAL INDIRECT
Depreciation
Taxes
Insurance
Other Fixed
TOTAL FIXED
TOTAL OPERATING
Credits
NET OPERATING
Used
Per NT
1.25 Ibs
0.02 NT
0.023 NT
-
8750 MH
-
570 FC
107.(L2+L3)
67 KWH
22.0 MG
MG
MCF
3.4 M Ib
-
-
30% L2
707o(L2+L3)
1000 MH
-
-
107. FC
270 FC
170 FC
-
-
-
-
Cost
Per Unit
$ .20
$40.
$30.
-
$3.50
$10,000
-
-
$.01
.05
.25
.30
.50
-
-
-
-
3.90
-
-
-
-
-
-
-
-
-
Per NT
Sulfur
$ .25
.80
.69
$ 1.74
.47
.08
2.95
.05
.67
1.10
-
1.70
3.47
$ 8.76
.14
-.38
-
.06
-
$ .58
5.91
1.18
.59
-
$ 7.68
$17.02
-
$17.02
Per Year
$M
$ 16.5
52.8
45.5
$ 114.8
31.0
5.0
195.0
3.6
44.2
72.5
-
112.0
228.7
$ 578.1
9.3
25.2
-
3.9
-
$ 38.4
390.0
78.0
39.0
-
$ 507.0
$1,123.5
-
$1,124.
5-27
-------
EXHIBIT 5-9
Capacity - Cost Relation
Per NT Sulfur
Case 2B (4.5% SOp) 200 NTD 100 NTD 400 NTD
Fixed Capital, MM $ (5.0) (3.3) (7.6)
Fixed Charges at 18% F.C. $ 13.65 $ 18.00 $ 10.39
R/M 3.08 3.08 3.08
Utilities 4.66 4.66 4.66
Direct Labor, Supv., Supplies .60 1.20 .30
Indirect Costs .58 1.16 .29
Total $ 22.57 $ 28.10 $ 18.72
Case 2C (2.9% S0;2) 200 NTD 100 NTD 400 NTD
Fixed Capital, ;MM $ (6.0) (4.0) (9.0)
Fixed Charges at 18% F.C. $ 16.38 $ 21.80 $12.30
R/M 4»88 4.88 4.88
Utilities 5.58 5.58 5.58
Direct Labor, Supv., Supplies .60 1.20 .30
Indirect Costs .58 1.16 .29
Total $ 28.02 $ 34.62 $ 23.35
Case 2D (8.0% SO..) 200 NTD 100 NTD 400 NTD
Fixed Capital, MM $ (3.9) (2.6) (5.9)
Fixed Charges at 18% F.C. $ 10.63 $ 14.20 $ 8.05
R/M - 1.74 1.74 1.74
Utilities 3.47 3.47 3.47
Direct Labor, Supv., Supplies .60 1.20 .30
Indirect Costs .58 1.16 .29
Total $ 17.02 $ 21.77 $ 13.85
5-28
-------
EXHIBIT 5-10
3
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5-29
-------
EXHIBIT 5-11
ECONOMI
SCb CONCENTRATI01
DMA SORPTION
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200 300 400 500
NTD Sulfur Equivalent in SO- Feed
6 7 8 9 -10
il 1W"
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600 700 800 900WO(
5-30
------- |