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A
PROJECT REPORT
ON
Production of Sulphuric Acid from DCDA Process
SUBMITTED BY:
AKSHAY AGARWAL (Roll No. 1104351003)
ANURAG VERMA (Roll No. 1104341009)
ISHA SHUKLA (Roll No. 1104351015)
Report Submission Date:
Submitted in
Partial fulfillment of the requirements for the awarding of
degree of
BACHELOR OF TECHNOLOGY IN
CHEMICAL ENGINEERING
Submitted To
UTTAR PRADESH TECHNICAL UNIVERSITY, LUCKNOW
UNDER THE EXPERT GUIDANCE OF:
Er. PRADEEP YADAV
DEPARTMENT OF CHEMICAL ENGINEERING
BUNDELKHAND INSTITUTE OF ENGINEERING AND TECHNOLOGY
JHANSI-284128
SESSION 2014-15
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CERTIFICATE
This is certified that Akshay Agarwal, Anurag Verma and Isha
Shukla have carried out this
project entitled Production of Sulphuric Acid from DCDA Process
for the award of
Bachelor of Technology from Uttar Pradesh Technical University,
Lucknow under my
supervision. The project embodies result of original work and
studies carried out by student
themselves and the contents of the project do not form the basis
for the award of any other
degree to the candidates or to anybody else.
Er. A.D. Hiwarikar Er. Pradeep Yadav
Head of the department Assistant Professor
BIET, Jhansi BIET, Jhansi
Date:
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ACKNOWLEDGEMENT
For the successful accomplishment of our project, we would like
to thank The Almighty for His
blessings. A special thanks to our project guide Er. Pradeep
Yadav whose help, stimulating
suggestions and encouragement, helped us to coordinate our
project especially in writing this
report. We rather find words short to express our gratitude to
him. His involvement and personal
interest has enabled us to accomplish this project work
successfully.
We are highly thankful to Er. A.D. Hiwarikar, Er. Sudeep Yadav,
Er.Ravindra Kumar , Er.
S.K.Srivastava, Er. Ajitesh Mehra and Er. Neeraj Singh,
Department of Chemical
Engineering, B.I.E.T. Jhansi for their full cooperation in
providing necessary facilities,
environment needed for the work.
Finally we wish to express our modest and sincere regards to our
parents and friends for their
intensive support and encouragement for this project work.
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ABSTRACT
The project includes the detailed designing of a four stage
adiabatic catalytic bed reactor for a
sulphuric acid plant of capacity 1000 TPD (tons per day). The
feed is at 1 atm and 4100 C. There
are two main processes for manufacturing of sulphuric acid
namely the chamber process and the
contact process. The pioneer sulphuric acid manufacturing
plants, adopted the chamber process
but at the beginning of the twentieth century with technological
advancements, the contact
process gained popularity as the conversion achieved was much
higher than that achieved
through chamber process. Chamber process produced sulphuric acid
of concentration less than
80 %.The major disadvantage includes the limitations in
throughput, quality and concentration of
the acid produced. All known new plants uses the contact process
although some older chamber
process plants may still be in use.
The contact process has been gradually modified to use double
absorption (also called double
catalyst), which increases yield and reduces stack emission of
unconverted SO2. Conversions
using single absorption contact process were typically about
97-98 percent. While in the current
double absorption flow process, achievable conversions are as
high as 99.7 percent.
The project mainly comprise of the basic parts of the sulphuric
acid manufacturing plant, the
equipments and the catalyst used, flow of materials in and out
of the equipments, their material
and energy balances, heat duty of the heat exchangers, weight of
the catalyst required and
pollution control.
.
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TABLE OF CONTENTS
1. List of Tables6
2. List of Figures...7
3. Notations...8
4. Introduction.....10
5. Literature Review....11
6. Uses and Applications.12
7. Sulphuric Acid- World Market....13
8. Status of Existing Sulphuric Acid Plants In India...14
9. The Contact Process........16
10. Available Technologies for Pollution Control........20
11. Material Balance........23
12. Energy Balance.......30
13. Weight of Catalyst......42
14. Summary Sheet.......55
15. Conclusion..57
16. References.......58
17. Appendix.....59
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1. LIST OF TABLES
Table 1: Capacity wise number of sulphuric acid plants in
India.
Table 2: Combustion Chamber Material Balance.
Table 3: Material Balance on First Three Catalytic Beds.
Table 4: Material Balance on Fourth Catalytic Bed.
Table 5: Heat Capacity equation constants for incoming gas
mixture.
Table 6: Heat Capacity equation constants for outgoing gas
mixture.
Table 7: Heat Capacity equation constants for incoming gas
mixture.
Table 8: Heat Capacity equation constants for outgoing gas
mixture.
Table 9: Heat Capacity equation constants for incoming gas
mixture.
Table 10: Heat Capacity equation constants for outgoing gas
mixture.
Table 11: Heat Capacity equation constants for incoming gas
mixture.
Table12: Heat Capacity equation constants for outgoing gas
mixture.
Table 13: Calculations of First Catalytic Bed.
Table 14: Rate Calculations of First Catalytic Bed.
Table 15: Calculations of Second Catalytic Bed.
Table 16: Rate Calculations of Second Catalytic Bed.
Table 17: Calculations of Third Catalytic Bed.
Table 18: Rate Calculations of Third Catalytic Bed.
Table 19: Calculations of Fourth Catalytic Bed.
Table 20: Rate Calculations of Fourth Catalytic Bed.
Table 21: Table for mole fractions expressed in terms of
conversion.
Table 22: Mole percent of gases entering the converter.
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2. LIST OF FIGURES
Figure 1: Structure of H2SO4 molecule.
Figure 2: Pie Chart showing Global Consumption of Sulphuric
Acid.
Figure 3: Sulfuric acid- structure of the world production
capacity by region, 2012.
Figure 4: Double Absorption Contact Process.
Figure 5: Four Stage Catalytic Reactor for Contact Process.
Figure 6: Activity of cesium based catalyst in comparison with
conventional catalyst.
Figure 7: Combustion Chamber Balance
Figure 8: Material balance over first three catalytic beds.
Figure 9: Primary absorption tower material balance.
Figure 10: Fourth catalytic bed material balance.
Figure 11: Final absorption tower material balance
Figure 12: Enthalpy balance over first catalytic bed.
Figure 14: Enthalpy balance over third catalytic bed.
Figure 15: Enthalpy balance over fourth catalytic bed
Figure 16: Plot of 1/-RA versus XA for first bed
Figure 17: Plot of 1/-RA versus XA for second bed
Figure 18: Plot of 1/-RA versus XA for third bed
Figure 19: Plot of 1/-RA versus XA for fourth bed
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3. NOTATIONS
HR Sum of enthalpies of all materials entering the reaction
process
relative to the reference state for the standard heat of
reaction at
298 K and 101.32 kPa.
H298 Standard heat of reaction at 298 K and 101.32 kPa.
q Net energy or heat added to the system.
HP Sum of enthalpies of all leaving materials referred to the
standard
reference state at 298 K
HRT Standard heat of reaction at temperature T (K)
[ (ni Hf) ]products Standard heat of formation of products.
[ (ni Hf )]reactants Standard heat of formation of
reactants.
T Temperature expressed in K
k1 Rate of forward reaction expressed in gmol/s-(gm
cat)-atm3/2
k2 Rate of backward reaction expressed in gmol/s-(gm
cat)-atm
NA Number of moles at conversion XA.
NA0 Initial number of moles.
Superficial mass velocity.
Fluid density
Superficial velocity in axial direction.
Reaction rate expressed in pseudo homogeneous form (i.e.
number
of moles transformed per unit time per unit of total reactor
volume)
Enthalpy change for the reaction at the indicated
conditions.
Bulk density of the catalyst (total mass of catalyst / total
volume of
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reactor)
=
Global reaction rate per unit mass of catalyst.
PA Partial pressure of A
Po Total pressure.
yA Mole fraction of A
Stoichiometric coefficient for reactant A (negative for
reactants)
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4. INTRODUCTION
Sulfuric acid is a highly corrosive strong mineral acid with the
molecular formula H2SO4 and
structure shown in the figure below:
Figure 1: Structure of H2SO4 molecule.
It is a pungent-ethereal, colorless to slightly yellow viscous
liquid which is soluble
in water at concentrations. Sometimes, it is dyed dark brown
during production to alert people to
its hazards. The historical name of this acid is oil of vitriol.
Sulphuric acid is an important
chemical, which has large-scale industrial uses. Its major user
is the phosphate fertilizer
industry. Other important applications are in petroleum
refining, steel pickling, rayon & staple
fiber, alum, explosives, detergents, plastics and fibers etc.
Sulphuric acid industry is very old and
has been continuously adopting the technological developments.
The progress made in sulphuric
acid manufacture during recent decades has led to changes in the
method and technology of its
manufacture, resulting mainly in the reduction of emissions of
sulphur compounds to air and
reduction of harmful waste. [3]
It started with Lead Chamber process followed by contact process
with Single
Conversion Single Absorption (SCSA) and now Double Conversion
Double Absorption Process
(DCDA). The Sulphuric Acid production through Contact Process is
very mature. However,
improvement in conversion and absorption stages are being
introduced from time to time to
increase conversion and absorption efficiencies, which also
result in reduction in emissions.
Most of the plants use elemental sulphur as raw material and in
few cases Copper/ Zinc Smelters
gases are being used to produce Sulphuric Acid. [3]
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5. LITERATURE REVIEW
Although sulphuric acid is now one of the most widely used
chemicals, it was probably little
known before the 16th century. It was prepared by Johann Van
Helmont (c.1600) by destructive
distillation of green vitriol (ferrous sulfate) and by burning
sulfur. [3]
In the seventeenth century, the German-Dutch chemist Johann
Glauber prepared sulfuric acid
by burning sulfur together with saltpeter (potassium nitrate,
KNO3), in the presence of steam. As
saltpeter decomposes, it oxidizes the sulfur to SO3, which
combines with water to produce
sulfuric acid. In 1736, Joshua Ward, a London pharmacist, used
this method to begin the first
large-scale production of sulfuric acid.[3]
In 1746 in Birmingham, John Roebuck adapted this method to
produce sulfuric acid in lead-
lined chambers, which were stronger, less expensive, and could
be made larger than the
previously used glass containers. Sulfuric acid created by John
Roebuck's process approached a
65% concentration. [3]
After several refinements, this method, developed into the
so-called the lead chamber
process or "chamber process. Later refinements to the lead
chamber process by French
chemist Joseph Louis Gay-Lussac and British chemist John Glover
improved concentration to
78%. However, the manufacture of some dyes and other chemical
processes require a more
concentrated product. Throughout the 18th century, this could
only be made by dry
distilling minerals in a technique similar to the original
alchemical processes. Pyrite (iron
disulfide, FeS2) was heated in air to yield iron(II) sulfate,
FeSO4, which was oxidized by further
heating in air to form iron(III) sulfate, Fe2(SO4)3, which, when
heated to 480 C, decomposed
to iron(III) oxide and sulfur trioxide, which could be passed
through water to yield sulfuric acid
in any concentration. However, the expense of this process
prevented the large-scale use of
concentrated sulfuric acid.[3]
In 1831, British vinegar merchant Peregrine Phillips patented
the contact process, which
was a far more economical process for producing sulfur trioxide
and concentrated sulfuric acid.
It was little used until a need for concentrated acid arose,
particularly for the manufacture of
synthetic organic dyes. Today, nearly all of the world's
sulfuric acid is produced using this
method. In the current flow process, achievable conversions are
as high as 99.7 percent.[3]
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6. USES AND APPLICATIONS
Sulfuric acid is one of the most important compounds made by the
chemical industry. It
is used to make, literally, hundreds of compounds needed by
almost every industry.
By far the largest amount of sulfuric acid is used to make
phosphoric acid, used, in turn,
to make the phosphate fertilizers. It is also used to make
ammonium sulfate, which is a
particularly important fertilizer in sulfur-deficient.
It is widely used in the manufacture of chemicals, e.g., in
making hydrochloric acid,
nitric acid, sulfate salts, synthetic detergents, dyes and
pigments, explosives, and drugs.
It is used in petroleum refining to wash impurities out of
gasoline and other refinery
products, and in waste water treatment.
Also widely used in metal processing for example in the
manufacture of copper and
the manufacture of zinc and in cleaning the surface of steel
sheet, known as pickling.
It is also used to make caprolactam, which is converted into
polyamide 6 and in
the manufacture of titanium dioxide, used, for example, as a
pigment.
It is used in the production of numerous goods including various
cleaning agents,
domestic acidic drain cleaners and electrolytes in lead-acid
batteries.[2]
Figure 2: Pie Chart showing Global Consumption of Sulphuric
Acid
Global Sulphuric Acid Consumption By End Use Sector in 2013
Fertilizer Production 56%
Other Applications 23%
Manufacture of Chemicals 11%
Agriultural Chemistry 10%
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7. SULPHURIC ACID - WORLD MARKET
The worldwide market for sulphuric acid witnessed stable growth
between 2009-2012,
supported by increasing demand from major end-use industries. In
2012, sulphuric acid
production grew by more than 7 million tonnes and exceeded 230.7
million tonnes. Asia
ranks as the leading sulfuric acid manufacturer, accounting for
around 45% of the overall
production. China, the US, India, Russia and Morocco are the top
five sulfuric acid
manufacturing countries.[6]
Figure 3: Sulfuric acid- structure of the world production
capacity by region, 2012
APAC (Uganda) is the major sulphuric acid consumer. In 2012, its
consumption volume
surpassed the 106 million mark. The fertilizer industry is the
products major end-use sector,
consuming over55% of the overall sulfuric acid output. In 2011,
the world foreign trade in
sulphuric acid was valued at more than USD (US dollar) 1.87
billion. Europe is the leading
sulphuric acid exporter, whilst Asia is a market leader in terms
of imports. The worldwide
sulphuric acid production is poised to increase in the
forthcoming years to go beyond 257.6
million by end-2015. [6]
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8. STATUS OF EXISTING SULPHURIC ACID PLANTS IN INDIA
In India, there are about 140 Sulphuric Acid Plants (130 Sulphur
based & 10 Smelter Gas
based) with Annual Installed Capacity of about 12 Million MT.
[6]
Table 1: Capacity wise number of sulphuric acid plants in
India
Installed Capacity (MT/Day) Number of Plants %
upto 50 18 12.9
51-100 45 32.1
101-200 40 28.6
201-300 17 12.1
301-500 5 3.6
501-1000 9 6.4
1001- 2000 4 2.9
above 2000 2 1.4
Total 140 100.0
The current annual production of Sulphuric Acid is about 5.5
Million MT, against the
installed capacity of 12 Million MT/Annum from Sulphur based as
well as Smelter Gas based
plants. The demand of Sulphuric Acid is fully met by the current
production, as the installed
capacity is more than double the demand.
The environmental problems arising due to Sulphuric Acid
manufacture include:
Off gases from absorption tower containing oxides of Sulphur
(SOx) and acid mist.
Liquid effluent generated through waste heat boiler blow-down,
spillage & leakage from
equipment, washing of equipment, cooling tower bleeding etc.
Generation of Solid Waste viz. Sulphur muck & spent
catalyst.
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Presently, Gaseous emission limits prescribed by CPCB for
Sulphuric Acid Plants are as under:
SO2 : 4.0 Kg/MT of Sulphuric Acid produced (Conc. 100%)
Acid Mist : 50 mg/Nm3
However, due to advancement in process and pollution control
technologies, it may be
possible to further reduce & control the emissions of SOx
and acid mist. In view of this, CPCB
took up a project to revisit the emission standards. [6]
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9. THE CONTACT PROCESS
Until 1990 there were no plants using the contact process, the
more popular process being
used was the chamber process. But, up to mid 1920s there were
considerable changes in the
contact process technology, equipments and the catalyst being
used. [2]
The primary raw material for producing intermediate product of
sulphur dioxide is
elemental sulphur. This can either be made by burning sulphur in
an excess of air or by heating
sulphide ores like pyrite in an excess of air. In either case,
an excess of air is used so that the
sulphur dioxide produced is already mixed with oxygen for the
next stage. [2]
These traditional contact plants can be further subdivided into
the double contact double
absorption process (DCDA), which is the type of process now most
commonly used in new
plants with intermediate absorption and the older, without
intermediate absorption also referred
to as the single contact single absorption process (SCSA). In
DCDA Plant, SO3 is removed from
the gas stream after 3rd bed, which shifts the equilibrium and
increases the rate of the forward
SO2 to SO3 reaction resulting in higher overall conversion and
reduces stack emission of
unconverted SO2.Conversions using single absorption contact
process were typically about 97-98
percent. While in the current double absorption flow process,
achievable conversions are as high
as 99.7 percent. [2]
6.1 Double Conversion Double Absorption (DCDA) Process
The main steps involved in DCDA process are as below:
Melting solid Sulphur with steam coils, followed by filtration
or settling of impurities to
obtain clean sulphur containing less than 10 mg/l of ash.
Burning the molten Sulphur with air to produce gas-containing
SO2.
Cooling the hot gas in Waste Heat Boiler System to produce
superheated or saturated steam
at conditions fixed, as per requirements.
Catalytic oxidation of SO2 to SO3 in three consecutive passes of
converter containing V2O5
catalyst with intercooling of gas in between. The exothermic
heat of reaction is utilized to
produce steam in Waste Heat Boiler system and to reheat the
gases going to pass IV from the
intermediate absorber.[6]
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The process flow sheet is shown below:
Figure 4: Double Absorption Contact Process
6.2 Chemical Reactions
In the manufacture of sulphuric acid from sulphur, the first
step is the burning of sulphur in a
furnace to form sulfur dioxide: [1]
S (l) + O2 (g) SO2 (g) H= -298.3 kJ
Following this step, the sulphur dioxide is converted to sulphur
trioxide, using a catalyst,
SO2 (g) + O2 (g) SO3 (g) H= -98.3 kJ
The final step is reacting sulphur trioxide with water to form
sulphuric acid.
SO3 (g) + H2O (l) H2SO4 (l) H= -130.4 kJ
6.3 Catalyst
A commercial sulphur dioxide- converting catalyst consists of
4-9 wt % vanadium pentaoxide,
V2O5, as the active component, together with alkali metal
sulphate promoters. At operating
temperatures the active ingredient is a molten salt held in a
porous silica pellet. Normally
V2O 5
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potassium sulphate is used as a promoter but in recent years
also caesium sulphate has been used.
Caesium sulphate lowers the melting point, which means that the
catalyst can be used at lower
temperatures. The carrier material is silica in different forms.
The lower temperature limit is 410-
430 C for conventional catalysts and 380-390 C for caesium doped
catalysts. The upper
temperature limit is 600-650 C above which, catalytic activity
can be lost permanently due to
reduction of the internal surface. These catalysts are long
lived up to twenty years and are not
subject to poisoning except fluorine. [2]
6.4 Contact Process Equipments
The main equipments being used in the process are
1) Combustion Chamber
2) Converters
3) Absorbers
4) Heat Exchangers
6.4.1 Converters - heart of the contact sulphuric acid plant
The reactor is often the central unit around which a chemical
plant is designed. Good
reactor design is thus important for the performance of the
plant. The chemical conversion of
sulphur dioxide to sulphur trioxide is designed to maximize the
conversion by taking into
consideration that:
1) Equilibrium is an inverse function of temperature and a
direct function of oxygen to sulphur
dioxide ratio.
2) Rate of reaction is a direct function of temperature.
3) Gas composition and amount of catalyst affect the rate of
conversion and kinetics of the
reaction.
4) Removal of sulphur trioxide formed allows more sulphur
dioxide to be converted.
The commercialization of these basic conditions makes possible
high overall conversion
by using a multi pass converter wherein, at an entering
temperature of 410C to 440C (the
ignition temperature), the major part of conversion takes place
(60 to 75 %) in the first catalytic
bed with an exit temperature of 600C or more, depending largely
on the concentration of
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sulphur dioxide in the gas. The successive lowering of
temperatures between stages ensures an
overall high conversion. [2]
The converter is usually provided with trays for supporting the
catalyst and manholes for
access to it. Converters have usually been made of cast iron and
aluminum-coated steel, but
stainless steel is now the preferred material of construction.
Pressure drop through the converter
must be minimized to reduce power consumption. All these must be
optimized to secure the
maximum yield and profit. [2]
6.5 Additional Data & Specification
Total Pressure = 1 atm
Feed composition (Mole Percent)
SO2 10
O2 11
N2 79
Overall Conversion = 99.8 %. [1]
The four stage catalytic reactor is shown below:
Figure 5: Four Stage Catalytic Reactor for Contact Process
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10. AVAILABLE TECHNOLOGIES FOR POLLUTION CONTROL
In upcoming plants, following features are considered
important:
Selection of 5-stage converter for maximum conversion
efficiency.
Use of sulphur filter for minimizing ash content.
Use of Cesium based catalyst in last bed of Converter for
maximum conversion efficiency.
Selection of high efficiency mist eliminators ensuring minimum
acid mist exhaust.
Use of Waste heat recovery from acid system.
Use of suitable start-up scrubbing system.[6]
10.1Modified Converter
Existing plants are generally based on 4 stage Converter except
for very few plants based on
5 stage Converter that has come up recently. 5-stage converter
helps in increasing the conversion
efficiency. This minimizes the stack emissions level of SO2.
With conventional catalyst the
conversion efficiency can be increased from 99.7% to 99.8%.
(1 kg SO2 instead of 2 Kg SO2 per MT of acid). [6]
10.2 Cesium Based Catalyst
The addition of Cesium Catalyst (CS) to the conventional
Alkali-Vanadium Catalyst has
long been known to enhance the low temperature properties of the
catalyst. Cesium based
Catalyst offers high activity at low operating temperature.
Emissions from existing plants can
roughly be cut in half without increasing catalyst volume. The
acid production capacity can be
increased by using higher strength sulphur dioxide gas without
increasing SO2 emissions and
plant pressure drop. New acid plants may be designed with low
SO2 emission by selecting
different type of Catalysts for different stages. [6]
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Figure 6: Activity of cesium based catalyst in comparison with
conventional catalyst.
10.3Mist Eliminators
Mist is inevitably formed at various points in Sulphuric Acid
Plants. If mist is unchecked,
and carried through rest of the plant in the gas stream, it
causes corrosion inside the plant and
environmental menace outside it. Mist is distinct from acid
spray, which is formed in the towers
by purely physical process of aspiration into the gas stream of
liquid droplets. Now a days, Mist
eliminators that are designed to remove virtually any type of
mist from any gas stream are
available. Mist eliminators excel at collecting, the very
difficult to remove sub micron size mist
particles from gas stream. [6]
10.4 Waste Heat and Heat Recovery System
The waste heat system is completely integrated in DCDA plants.
In economizers that cool
the gases from third & fourth bed of converter, heat is
utilized for preheating feed water for
WHB System. Heat generated in Sulphur furnace, heats up this
feed water and steam is
generated at about 2500C temperatures. This steam is superheated
to about 4000 C for cooling the
1st stage out converter gases. This superheated steam can be
used for generating power and
saturated steam for process heating. The Heat Recovery System is
basically an absorber that
operates at about 2000 C and uses a boiler to remove the
absorption heat as low pressure steam
(at upto 10 bar), instead of acid coolers (where heat is
wasted). [6]
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10.5 Scrubbing System
In DCDA Sulphuric Acid plants, emission levels of SO2 are higher
during start-up or shut down
when SO2 to SO3 conversion is not proper. Also in SCSA Plant,
the emission levels of SO2 are
normally high. While in first case, the start-up scrubbers are
required to take care of extra SO2
load during unstabilized conditions, start-up & shutdown, in
second case, a continuous scrubbing
unit is required to take care of tail gases going out for the
stack.[7]
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11. MATERIAL BALANCE
11.1 Assumptions
1. Complete burning of S in the burner.
2. 99.8% conversion of SO2 toSO3 in the reactor.
3. Overall absorption of SO3 in the process is 100%
4. 40% excess oxygen is provided.
5. Humidity of entering air is 65% at 300C
11.2 Calculations
Capacity: 1000 TPD H2SO4 plant
Basis: 1 hr of operation
Purity: 98 % pure acid
Reactions:
S (l) + O2 (g) SO2 (g)
SO2 (g) + O2 (g) SO3 (g)
SO3 (g) + H2O (l) H2SO4 (l)
1000 TPD H2SO4 = (1000 x 103)/ 24 = 41,666.67 kg/hr
98% pure acid produced = 41,666.67 x 0.98 = 40,833 kg/hr
No. of moles of acid produced = 40,833.33 / 98 = 416.67
kmol/hr
(Molecular weight of H2SO4=98)
Overall absorption of acid = 100 %
SO3 (g) + H2O (l) H2SO4 (l)
Therefore, by stoichiometry,
SO3 required =416.67 kmol/hr
Overall conversion of SO2 to SO3 = 99.8 %
V2O 5
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SO2 (g) + O2 (g) SO3 (g)
Let SO2formed be X kmol/hr
0.998 X=416.67
SO2 required =X= 416.67/0.998 = 417.501 kmol/hr
O2 required = 417.501 x 0.5 =208.75 kmol/hr
11.1.1 Combustion Chamber Balance
Figure 7: Combustion Chamber Balance
S (l) + O2 (g) SO2 (g) (Assuming 100 % combustion)
S required = 417.501 kmol/hr =13,360.053 kg/hr
O2 required = 417.501 x 1 = 417.501 kmol/hr
Total O2 required = 417.501 + 208.75 = 626.25 kmol/hr
O2 is taken in 40% excess
O2 in the combustion chamber = 626.25 x 1.4 = 876.75 kmol/hr
(Dry air contains 21% O2)
Dry air in = 876.765/0.21 = 4175.07 kmol/hr
(Molecular weight of air=29)
Dry air in = 4175.07 x 29 = 121076.381 kg/hr
V2O 5
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Table 2: Combustion Chamber Material Balance
Component Inlet (kmol) Inlet ( kg) Outlet (kmol) Outlet (kg)
S
O2
N2
SO2
Total
417.501
876.75
3298.306
-
4592.557
13,360.032
28,056.00
92,352.568
-
133,768.6
-
459.249
3298.306
417.501
4175.056
-
14,695.968
92,352.568
26,720.064
133,768.6
11.2.2 Overall Balance over First Three Catalytic Beds
Figure 8: Material balance over first three catalytic beds
SO2 (g) + O2 (g) SO3 (g) (Conversion =96.7 %)
SO2 in = 417.501 kmol
O2 in = 459.249 kmol
SO2 reacted = 417.501*0.967=403.723 kmol/hr
O2reacted = 0.5*403.723 = 201.86 kmol/hr
SO3formed = 403.723 kmol/hr
SO2out = SO2 in SO2reacted
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=417.501 403.723 =13.778 kmol/hr
O2out = O2 in O2reacted
= 459.249 201.86 =257.389 kmol/hr
Table 3: Material Balance on First Three Catalytic Beds
Component Inlet (kmol) Mole % Inlet (kg) Outlet (kmol) Outlet
(kg)
SO2
O2
N2
SO3
Total
417.501
459.249
3298.306
-
4,175.056
10 26,720.064 13.778 881.792
11 14,695.968 257.389 8236.448
79 92,352.568 3298.306 92,353.568
- - 403.723 32297.84
100 133,768.6 3973.196 133769.648
After the third stage, 40 % of product goes to economizer and
then to the Interpass Absorber.
11.2.3 Primary Absorber
Figure 9: Primary absorption tower material balance
-
27
(Assuming 100 % absorption)
SO2 in = 0.4*13.778 = 5.5112 kmol/hr
SO3 in = 0.4*403.723 = 161.489 kmol/hr
O2 in = 0.4 * 257.389 = 102.955 kmol/hr
N2 in = 0.4 * 3298.306 = 1319.3224 kmol/hr
SO3 (g) + H2O (l) H2SO4 (l)
H2SO4 formed = 161.489 kmol/hr
(Sulphur dioxide, Oxygen and Nitrogen are recycled from inter
pass absorber to the 4th stage of
reactor).
11.2.4 Fourth Catalytic Bed
Figure 10: Fourth catalytic bed material balance
Conversion = 3.1 % (Overall = 99.8 %)
Gases in to the fourth stage of reactor constitute 60 % of
product from third stage and gases
recycled from Interpass absorber.
SO2 in = Total SO2out from the 3rd bed = 13.778 kmol/hr
O2 in = Total O2out from the 3rd bed = 257.389 kmol/hr
N2 in = 3298.306 kmol/hr
-
28
SO3 in = 60% of SO3out from the 3rd bed
= 0.6 * 403.723 = 242.23 kmol/hr
SO2 (g) + O2 (g) SO3 (g) (Overall Conversion =99.8 %)
SO2 reacted upto 4th bed = 99.8 % of SO2 entering the 1st
bed
= 0.998 * 417.501 = 416.67 kmol/hr
SO2 out = Initial SO2 Total SO2 reacted
= 417.501 416.67 = 0.831 kmol/hr
Therefore SO2 reacted in the 4th bed
= SO2 in SO2 out =13.778 0.831 = 12.947 kmol/hr
O2 reacted = 12.947 * 0.5= 6.473 kmol/hr
SO3 formed = 12.947 kmol/hr
O2 out = O2 in O2 reacted
=257.389 6.473 = 250.916 kmol/hr
Total SO3 outlet = SO3 inlet from the 3rd bed+ SO3 formed in the
4th bed
= 242.23 + 12.947 = 255.177 kmol/hr
Table 4: Material Balance on Fourth Catalytic Bed
Component Inlet (kmol) Inlet (kg) Outlet (kmol) Outlet (kg)
SO2 13.778 881.792 0.831 53.184
O2 257.389 8,236.448 250.916 8,029.312
N2 3298.306 92,353.568 3298.306 92,353.568
SO3 242.23 19378.4 255.177 20,414.16
Total 3,809.703 120,850.208 3,805.23 120,850.224
-
29
11.2.5 Final Absorption Tower
Figure 11: Final absorption tower material balance
SO3 in = 255.173 kmol/hr
SO2 in = 0.831 kmol/hr
O2 in = 250.916 kmol/hr
N2 in = 3298.306 kmol/hr
H2SO4 formed = 255.173kmol/hr
Total H2SO4 formed = 255.173 + 161.489=416.662kmol/hr
-
30
12. ENERGY BALANCE
12.1Theory
12.1.1 Law of Conservation of Energy
Energy can neither be created nor be destroyed; it can only be
changed from one form to
another. In simpler words, all energy entering a process is
equal to that leaving plus that left in
the process. Energy can appear in many forms. Some of the forms
are enthalpy, electrical energy,
chemical energy (in terms of H reaction), kinetic energy,
potential energy, work, and heat
inflow.
In many cases in process engineering , which often takes place
at constant pressure ,
electrical energy , kinetic energy, potential energy, and work
either are not present or can be
neglected. Then only the enthalpy of the materials (at constant
pressure), the standard chemical
reaction energy (H) at 25 C, and the heat added or removed must
be taken into account in the
energy balance .This is called heat balance. [5]
12.1.2 Heat Balance
The energy or heat coming into a process in the inlet materials
plus any net energy added to
the process is equal to the energy leaving in the materials.
Expressed mathematically,
HR + ( H298 ) + q = HP (1)
where, HR is the sum of enthalpies of all materials entering the
reaction process relative to the
reference state for the standard heat of reaction at 298 K and
101.32 kPa.
H298 = standard heat of reaction at 298 K and 101.32 kPa.
The reaction contributes heat to the process, so the negative of
H298 is taken to be positive
input heat for an exothermic reaction. [5]
q = net energy or heat added to the system.
HP =sum of enthalpies of all leaving materials referred to the
standard reference state at 298 K.
-
31
12.2 Equations for calculating net enthalpy change at catalyst
bed[2]
The temperature dependency of the heat capacity is given by the
equation
CP = a+bT+cT2 + (2)
HRT = T 298 (ni Cpi) Reactants dT + H298 + 298 T (ni Cpi)
Products dT (3)
where, HRT is the standard heat of reaction at temperature T
(K)
HRT = H298 + 298T ( (ni Cpi) Products (ni Cpi) Reactants)dT
(4)
Enthalpy change, with 298 K as the reference temperature, can be
calculated from the formula,
H = HRT H298 = ( ai ni )(T298) + (( bi ni)/2)(T2 2982) + (5)
Similarly, Enthalpy change between T1 K and T2 K, can be
calculated from the formula
H = ( ai ni )(T1T2) + (( bi ni)/2)(T12 T22) + (6)
Standard heat of reaction at 298 K (H298 ) from heat of
formation is given by
H298 = [ (ni Hf) ]Products [ (ni Hf )]Reactants (7)
where [ (ni Hf) ]products = standard heat of formation of
products and
[ (ni Hf )]reactants = standard heat of formation of
reactants
Net heat to be removed from the catalytic bed ,
q = HP HR + H298 (8)
12.3 Data given
12.3.1 Inlet and Effluent Temperatures
First stage: Inlet temperature = 410C
Effluent temperature = 602C
Second stage: Inlet temperature = 438C
Effluent temperature = 498C
Third stage: Inlet temperature = 432C
Effluent temperature = 444 C
-
32
Fourth stage: Inlet temperature = 427C
Effluent temperature = 435C
12.3.2 Heat Capacities[1]
N2= 29.5909 5.14 x 10-3T (9)
O 2=26.0257 + 11.755 x 10-3T (10)
SO 2 = 24.7706 + 62.9481 x 10-3 T (11)
SO 3 = 22.0376 + 121.624 x 10-3 T (12)
where , Heat capacity is expressed in kJ/kmol-K and T is
temperature in K.
12.4 Calculations
a = 22.036 24.771 0.5(26.026) = 15.748
b*103 = 121.624 62.948 0.5(11.755) = 52.799
c*106 = 91.867 ( 44.258) 0.5(2.343) = 46.438
d*109 = 24.369 11.122 0.5( 0.562) = 13.258
From equation (7)
H298 = 395720 ( 296810) = 98910 kJ/kmol SO2
Substituting T=298 K in equation (6)
98910 = Ho + ( 15.748*298) + (((52.799 * 103 ) /2)*2982) +
(((46.438 * 106 )/3)*2983)
+ (((13.258 * 109 )/4)*2984)
Ho = 96178kJ/kmol SO2 reacted
Substituting the value of Ho in equation (6), we get
HRT = 96178 15.748 T+ (26.4*10-3) T2 (15.48*10-6)T3+
(3.382*10-9)T4 (13)
-
33
12.4.1 Calculation of Heat Duty Required In Each Bed
First Catalytic Bed
Table 5: Heat Capacity equation constants for incoming gas
mixture
Component ni (kmol/hr.) ai aini bi bini*103
SO2 417.501 24.7706 10341.75 62.9481 26280.89
O2 459.249 26.0257 11952.2767 11.7551 5398.518
N2 3298.306 29.5909 97599.843 -5.141 -16956.59
SO3 - 22.0376 - 121.624 -
Total 4175.056 - 119893.8697 - 14722.818
Enthalpy of incoming gas mixture at 683K over 298K can be
calculated by substituting T=683 in
equation (5),
HR = 119893.8697(683298) + ((14722.818*10-3)/2) (68322982)
=48939432.44 kJ/hr
HR = (48939432.44 /3600) = 13594.286 kW
Table 6: Heat Capacity equation constants for outgoing gas
mixture
Component ni (kmol/hr.) ai aini bi bini*103
SO2 108.55 24.7706 2688.848 62.9481 6833.016
O2 304.77 26.0257 7931.852 11.7551 3582.6018
N2 3298.306 29.5909 97599.843 -5.141 -16956.59
SO3 308.95 22.0376 6808.516 121.624 37575.7348
Total 4020.576 - 115029.0566 - 31034.76286
Enthalpy of outgoing gas mixture at 875K over 298K can be
calculated by substituting T=875 in
equation (5),
Hp = 115029.0566(875298) + ((31034.763*10-3)/2)(87522982)
=76874549.2 kJ/hr
Hp = (76874549.2 /3600) = 21354.04144 kW
-
34
Total heat of reaction at 298 K,
H298 = 98910 kJ/kmol SO2 reacted = (417.501 108.55) (98910/3600)
= 8488.428 kW
Net enthalpy change, q = HP HR + H298
q = 21354.04144 13594.286 8488.428 = 728.67 kW
Thus, during the reaction, enthalpy equivalent to -728.67 kW is
to be removed from the first
catalytic bed in order to maintain the temperature of the
outgoing gas mixture at 875K.
Figure 12: Enthalpy balance over first catalytic bed
Second Catalytic Bed
Table 7: Heat Capacity equation constants for incoming gas
mixture
Component ni (kmol/hr.) ai aini bi bini*103
SO2 108.55 24.7706 2688.848 62.9481 6833.016
O2 304.77 26.0257 7931.852 11.7551 3582.6018
N2 3298.306 29.5909 97599.843 -5.141 -16956.59
SO3 308.95 22.0376 6808.516 121.624 37575.7348
Total 4020.576 - 115029.0566 - 31034.76286
-
35
Enthalpy of incoming gas mixture at 711K over 298K can be
calculated by substituting T=711 in
equation (5),
HR = 115029.056(711298) + ((31034.763*10-3)/2) (71122982)
=53973356.79 kJ/hr
HR = (53973356.79 /3600) = 14992.599 kW
Table 8: Heat Capacity equation constants for outgoing gas
mixture
Component ni (kmol/hr.) ai aini bi bini*103
SO2 31.73 24.7706 785.97 62.9481 1997.343
O2 266.36 26.0257 6932.2054 11.7551 3131.088
N2 3298.306 29.5909 97599.843 -5.141 -16956.59
SO3 385.77 22.0376 8501.144 121.624 46918.89
Total 3982.166 - 113819.46 - 35090.7314
Enthalpy of outgoing gas mixture at 778K over 298K can be
calculated by substituting T=778 in
equation (5),
Hp = 1153819.46(778298) + ((35090.7314*10-3)/2)(77822982)
=60879790.44 kJ/hr
Hp = (60879790.44 /3600) = 16911.053 kW
Total heat of reaction at 298 K,
H298 = 98910 kJ/kmol SO2 reacted
= (108.5531.73) (98910/3600) = 2110.6295 kW
Net enthalpy change, q = HP HR + H298
q = 16911.053 14992.599 2110.6295 = 192.1755 kW
Thus, during the reaction, enthalpy equivalent to 192.1755 kW is
to be removed from the
second catalytic bed in order to maintain the temperature of the
outgoing gas mixture at 778 K.
-
36
Figure 13: Enthalpy balance over second catalytic bed
Third Catalytic Bed
Table 9: Heat Capacity equation constants for incoming gas
mixture
Component ni (kmol/hr.) ai aini bi bini*103
SO2 31.73 24.7706 785.97 62.9481 1997.343
O2 266.36 26.0257 6932.2054 11.7551 3131.088
N2 3298.306 29.5909 97599.843 -5.141 -16956.59
SO3 385.77 22.0376 8501.144 121.624 46918.89
Total 3982.166 - 113819.46 - 35090.7314
Enthalpy of incoming gas mixture at 705K over 298K can be
calculated by substituting T=705 in
equation (5),
HR = 113819.46(705298) + ((35090.7314*10-3)/2) (70522982)
=53486906.95 kJ/hr
HR = (53486906.95 /3600) = 14857.474 kW
-
37
Table 10: Heat Capacity equation constants for outgoing gas
mixture
Component ni (kmol/hr.) ai aini bi bini*103
SO2 13.77 24.7706 341.091 62.9481 866.79
O2 257.38 26.0257 6698.494 11.7551 3025.5276
N2 3298.306 29.5909 97599.843 -5.141 -16956.59
SO3 403.77 22.0376 8898.1217 121.624 49103.257
Total 39732.226 - 113536.669 - 36038.989
Enthalpy of outgoing gas mixture at 717K over 298K can be
calculated by substituting T=717 in
equation (5),
Hp = 113536.669(717298) + ((36038.989*10-3)/2) (71722982)
=55235285.03 kJ/hr
Hp = (55235285.03 /3600) = 15343.1347 kW
Total heat of reaction at 298 K,
H298 = 98910 kJ/kmol SO2 reacted
= (31.7313.77) (98910/3600) = 493.451 kW
Net enthalpy change, q = HP HR + H298
q = 15343.1347 14857.474 493.451= 7.7903 kW
Thus, during the reaction, enthalpy equivalent to 7.7903 kW is
to be removed from the third
catalytic bed in order to maintain the temperature of the
outgoing gas mixture at 717 K.
-
38
Figure 14: Enthalpy balance over third catalytic bed
Fourth Catalytic Bed
Table 11: Heat Capacity equation constants for incoming gas
mixture
Component ni (kmol/hr.) ai aini bi bini*103
SO2 13.77 24.7706 341.091 62.9481 866.79
O2 257.38 26.0257 6698.494 11.7551 3025.5276
N2 3298.306 29.5909 97599.843 -5.141 -16956.59
SO3 242.238 22.0376 5338.344 121.624 29461.9545
Total 39732.226 - 109977.772 - 16397.686
.
Enthalpy of incoming gas mixture at 700K over 298K can be
calculated by substituting T=700 in
equation (5),
HR = 109977.772(700298) + ((16397.686*10-3)/2) (70022982)
=47500407.36 kJ/hr
HR = (47500407.36 /3600) = 13194.5576 kW
-
39
Table12: Heat Capacity equation constants for outgoing gas
mixture.
Component ni (kmol/hr.) ai aini bi bini*103
SO2 0.835 24.7706 20.6834 62.9481 52.5616
O2 250.9125 26.0257 6530.1734 11.7551 2949.5015
N2 3298.306 29.5909 97599.843 -5.141 -16956.59
SO3 255.173 22.0376 5623.4 121.624 31035.1609
Total 3805.2265 - 109774.0998 - 17080.634
Enthalpy of outgoing gas mixture at 708K over 298K can be
calculated by substituting T=708 in
equation (5),
Hp = 109775.0998(708298) + ((17080.634*10-3)/2)(70822982)
=48773723 kJ/hr
Hp = (48773723/3600) = 13548.256 kW
Total heat of reaction at 298 K,
H298 = 98910 kJ/kmol SO2 reacted
= (13.770.835) (98910/3600) = 355.389 kW
Net enthalpy change, q = HP HR + H298
q = 13480.533 13194.5576 355.389= 1.69 kW
Thus, during the reaction, enthalpy equivalent to 1.69 kW is to
be removed from the fourth
catalytic bed in order to maintain the temperature of the
outgoing gas mixture at 708 K.
-
40
Figure 15: Enthalpy balance over fourth catalytic bed
12.4.2 Calculation of Heat Load for Heat Exchangers
Heat load required between first and second catalytic bed:
At the outlet of first catalytic bed:
aini = 115029.0566
bini = 31034.76286* 10-3
First stage outlet temperature = 602C = 875K
Second stage inlet temperature = 438C =711K
Temperature change = 875 K to 711 K
Heat load is calculated by substituting T1=875K and T2=711K in
equation (6),
H1 = 115029.056(711875) + ((31034.763*10-3)/2) (71128752) =
22900898.26 kJ/hr
H1 = (22900898.26 /3600) = 6361.36 kW
Heat load required between second and third catalytic bed:
At the outlet of second catalytic bed:
aini = 113819.46
bini = 35090.7314* 10-3
-
41
Second stage outlet temperature = 498C = 771K
Third stage inlet temperature = 432C =705K
Temperature change = 771 K to 705 K
Heat load is calculated by substituting T1=711K and T2=705K in
equation (6),
H2 = 113819.46(705711) + ((35090.7314*10-3)/2) (70527112) =
831982.187 kJ/hr
H2 = (831982.187/3600) = 231.106 kW
Heat load required between third and fourth catalytic bed:
At the outlet of third catalytic bed:
aini = 109977.772
bini = 16397.686* 10-3
Third stage outlet temperature = 444 C = 717 K
Fourth stage inlet temperature = 427C =700 K
Temperature change = 717 K to 700 K
Heat load is calculated by substituting T1=700K and T2=717K in
equation (6),
H3 = 109977.772(700717) + ((16397.686*10-3)/2) (70027172) =
1996876.366 kJ/hr
H3 = (1996876.366 /3600) = 544.68 kW
-
42
13. WEIGHT OF CATALYST
13.1 Theory and Data[1]
The kinetic data are represented by a rate expression of the
form
=122 232
0.5
20.5 (14)
that may be regarded as a degenerate form of typical
Hougen-Watson kinetics. The rate constants
are given by
ln 1 = 12.07 31000 (15)
ln 2 = 22.75 53600 (16)
where
T is temperature expressed in K
R is expressed in cal/gmolK
k1 is expressed in gmol/s-(gm cat)-atm3/2
k2 is expressed in gmol/s-(gm cat)-atm
Derived equation for relationship between temperature and
conversion:
= + 307.158( )
1 + 0.0248( ) (17)
Where, T is reference temperature and XA is conversion at this
temperature.
T0 is initial temperature and XA0 is conversion at initial
conditions.
Relationship between number of moles and conversion:[4]
= (1 )
(1 + )
(18)
where
NA is number of moles at conversion XA.
NA0 is initial number of moles.
-
43
T0/T is temperature correction factor.
Relationship between partial pressure and mole fraction:
PA = yA Po (19)
PA is partial pressure of A and Po is total pressure.
yA is mole fraction of A.
The partial pressures of the various species are numerically
equal to their mole fractions since
the total pressure is one atmosphere.
Equation of Plug flow reactor for calculating weight of
catalyst:[4]
=
(20)
where,
W is weight of catalyst required.
FA0 is molar flow rate of Sulphur dioxide entering in feed.
13.2 Calculations
13.2.1 First Catalytic Bed
Sample Calculation:
For T = 773 K, T0 = 683 K, XSO2 = 0
773 =683 + 307.158 2
1 + 0.0248 2
2 = 0.3125
2 = 417.501(1 0.3125)
(1 0.05 0.3125)
683
773
2 = 257.633
2 = 417.501 257.633 = 159.868
3 = 159.868
-
44
2 = 459.249 (159.868 2 ) = 379.315
= 257.633 + 159.868 + 379.315 + 3298.306 = 4095.122
2 = 2 =159.868
4095.122= 0.062
2 = 2 =379.315
4095.122= 0.0926
3 = 3 =159.868
4095.122= 0.0390
Table 13: Calculations of First Catalytic Bed
T(K) XSO2 NSO2 NO2 NSO3 NN2 TOTAL PSO2(atm) PO2(atm)
PSO3(atm)
683 0 417.501 459.249 0 3298.306 4175.056 0.099 0.1099 0
688 0.0172 407.674 454.335 9.826 3298.306 4170.143 0.097 0.108
0.0023
725 0.1452 338.649 419.823 78.851 3298.306 4135.63 0.081 0.1015
0.0190
748 0.2252 298.726 399.861 118.77 3298.306 4115.669 0.072 0.0971
0.0288
773 0.3125 257.633 379.315 159.86 3298.306 4095.122 0.062 0.0926
0.0390
793 0.382 226.33 363.664 191.16 3298.306 4079.472 0.055 0.0891
0.0468
798 0.400 218.711 359.854 198.78 3298.306 4075.661 0.053 0.0882
0.0487
823 0.488 181.753 341.375 235.74 3298.306 4057.182 0.044 0.0841
0.0581
848 0.576 146.579 323.788 270.92 3298.306 4039.595 0.036 0.0801
0.0670
856.4 0.606 135.132 318.064 282.36 3298.306 4033.871 0.033
0.0788 0.07
873 0.665 113.026 307.011 304.47 3298.306 4022.819 0.028 0.0763
0.0756
875 0.6726 110.408 305.702 307.09 3298.306 4021.51 0.027 0.0760
0.0763
-
45
Sample Calculation:
1 = (12.07 31000
(1.987 773)) = 3 104
2 = (22.75 536000
(1.987 773)) = 5.3 106
=((3 104) 0.062 0.0926) ((5.3 106) 0.0390 0.09260.5)
0.0620.5
= 6.71 106
1 = 149047.3
Table 14: Rate Calculations of First Catalytic Bed
k1 k2 (-rA) (kmol/s-kg catalyst) 1/(-rA)
2.09E-05 5.34E-08 7.27E-07 1375530
2.47E-05 7.12E-08 8.41E-07 1188741
7.87E-05 5.26E-07 2.27E-06 439566
1.53E-04 1.65E-06 3.94E-06 253862.4
3.00E-04 5.30E-06 6.71E-06 149047.3
4.98E-04 1.28E-05 9.7E-06 103124.2
5.64E-04 1.58E-05 1.05E-05 94832.17
1.02E-03 4.42E-05 1.47E-05 68182
1.79E-03 1.16E-04 1.57E-05 63794.88
2.14E-03 1.59E-04 1.38E-05 72398.6
3.02E-03 2.88E-04 2.76E-05 362480.6
3.15E-03 3.10E-04 2.98E-05 378674.7
-
46
Figure 16: Plot of 1/-RA V/S XA for first bed
FA0 = 417.501 kmol/hr = 417.501/3600 kmol/s =
0.1159725kmol/s
W = FA0 * Area under the curve= 0.1159725*250920
Weight of catalyst in first catalyst bed =29103.3 kg
13.2.2 Second Catalytic Bed
Sample Calculation:
For T = 715 K, T0 = 711 K, 2 = 0.672
715 =711 + 307.158 (2 0.672)
1 + 0.0248 (2 0.672)
2 = 0.686
2 = 417.501(1 0.686)
(1 0.05 0.686)
683
715
2 = 129.523
2 = 417.501 129.523 = 287.978
3 = 287.978
2 = 459.249 (287.978 2 ) = 315.26
-
47
= 129.523 + 287.978 + 315.26 + 3298.306 = 4031.067
2 = 2 =129.523
4031.067= 0.03213
2 = 2 =315.26
4031.067= 0.0782
3 = 3 =287.978
4031.067= 0.0714
Table 15: Calculations of Second Catalytic Bed
T(K) XSO2 NSO2 NO2 NSO3 NN2 TOTAL PSO2(atm.) PO2(atm)
PSO3(atm)
711 0.672 110.408 305.702 307.09 3298.306 4021.51 0.027454
0.076017 0.076363
715 0.686 129.523 315.260 287.97 3298.306 4031.067 0.032131
0.078208 0.07144
718 0.696 124.784 312.890 292.71 3298.306 4028.697 0.030974
0.077665 0.072658
723 0.714 116.957 308.977 300.54 3298.306 4024.784 0.029059
0.076769 0.074673
728 0.731 109.22 305.108 308.28 3298.306 4020.915 0.027163
0.07588 0.076669
733 0.748 101.57 301.283 315.93 3298.306 4017.09 0.025284 0.075
0.078647
738 0.766 94.005 297.501 323.49 3298.306 4013.308 0.023423
0.074129 0.080606
743 0.783 86.523 293.760 330.97 3298.306 4009.567 0.021579
0.073265 0.082547
748 0.800 79.124 290.060 338.37 3298.306 4005.868 0.019752
0.072409 0.08447
753 0.818 71.805 286.401 345.69 3298.306 4002.208 0.017941
0.071561 0.086376
758.3 0.836 64.132 282.564 353.36 3298.306 3998.372 0.01604
0.07067 0.088378
763 0.85 57.4009 279.19 360.10 3298.306 3995.006 0.014368
0.069887 0.090138
768 0.87 50.312 275.654 367.18 3298.306 3991.462 0.012605
0.069061 0.091994
771 0.88 46.094 273.545 371.40 3298.306 3989.353 0.011554
0.068569 0.093099
778 0.90 36.355 268.676 381.14 3298.306 3984.483 0.009124
0.067431 0.095658
-
48
Table 16: Rate Calculations of Second Catalytic Bed
k1 k2 (-rA) (kmol/s-kg catalyst) 1/(-rA)
5.16E-05 2.53E-07 6.11585E-07 1635095
5.83E-05 3.13E-07 7.8216E-07 1278512
6.38E-05 3.66E-07 8.36552E-07 1195383
7.42E-05 4.75E-07 9.34367E-07 1070243
8.61E-05 6.14E-07 1.0383E-06 963113
9.96E-05 7.89E-07 1.14852E-06 870687.1
1.15E-04 1.01E-06 1.26386E-06 791225.4
1.33E-04 1.30E-06 1.38501E-06 722018.9
1.53E-04 1.65E-06 1.28608E-06 777554.9
1.75E-04 2.10E-06 1.31821E-06 758606.4
2.03E-04 2.69E-06 1.3143E-06 760864
2.29E-04 3.36E-06 1.25102E-06 799348
2.63E-04 4.42E-06 1.08457E-06 922027.9
2.84E-04 4.85E-06 9.94423E-07 1005608
3.41E-04 6.94E-06 3.92238E-07 2549474
Sample Calculation:
1 = (12.07 31000
(1.987 715)) = 5.828 105
2 = (22.75 536000
(1.987 715)) = 3.1278 107
=((5.828 105) 0.03213 0.0782) ((3.1278 107) 0.0714
0.07820.5)
0.032130.5
= 7.8216 107
-
49
1 = 1278512
Figure 17: Plot of 1/-RA v/s XA for second bed
FA0 = 110.408kmol/hr = 110.408/3600 kmol/s= 0.036688 kmol/s
W = FA0 * Area under the curve = 0.036688*161430
Weight of catalyst in first catalyst bed = 4950.878 kg
-
50
13.2.3 Third Catalytic Bed
Table 17: Calculations of Third Catalytic Bed
T(K) XSO2 NSO2 NO2 NSO3 NN2 TOTAL PSO2(atm) PO2(atm)
PSO3(atm)
705 0.905 36.355 268.67 381.14 3298.306 3984.483 0.00912 0.0674
0.09565
707 0.912 37.152 269.07 380.34 3298.306 3984.882 0.00932 0.0675
0.09544
708 0.915 35.648 268.32 381.85 3298.306 3984.13 0.00894 0.0673
0.09584
710 0.922 32.652 266.82 384.84 3298.306 3982.632 0.00819 0.0669
0.09663
712 0.929 29.669 265.33 387.83 3298.306 3981.14 0.00745 0.0666
0.09741
714 0.936 26.700 263.84 390.80 3298.306 3979.656 0.00670 0.0662
0.09819
715 0.939 25.221 263.10 392.27 3298.306 3978.916 0.00633 0.0661
0.09859
716 0.943 23.745 262.37 393.75 3298.306 3978.178 0.00596 0.0659
0.09897
717 0.946 22.272 261.63 395.22 3298.306 3977.442 0.0056 0.0657
0.09936
Sample Calculation:
For T = 708 K, T0 = 705 K, 2 = 0.905
708 =705 + 307.158 (2 0.905)
1 + 0.0248 (2 0.905)
2 = 0.915
2 = 417.501(1 0.915)
(1 0.05 0.915)
683
708
2 = 35.648
2 = 417.501 35.648 = 381.85
3 = 381.85
2 = 459.249 (381.85 2 ) = 268.32
= 35.648 + 268.32 + 381.85 + 3298.306 = 3984.13
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51
2 = 2 =35.648
3984.13= 0.00894
2 = 2 =268.32
3984.13= 0.0673
3 = 3 =381.85
3984.13= 0.09584
Table 18: Rate Calculations of Third Catalytic Bed
k1 k2 (-rA) (kmol/s-kg catalyst) 1/(-rA)
4.27E-05 1.83E-07 2.27E-07 4396699
4.55E-05 2.04E-07 2.44E-07 4094072
4.69E-05 2.15E-07 2.42E-07 4129794
4.99E-05 2.39E-07 2.37E-07 4224889
5.32E-05 2.67E-07 2.28E-07 4383353
5.65E-05 2.96E-07 2.15E-07 4641287
5.83E-05 3.13E-07 2.07E-07 4825010
6.01E-05 3.29E-07 1.98E-07 5053386
6.19E-05 3.47E-07 1.87E-07 5361426
Sample Calculation:
1 = (12.07 31000
(1.987 708)) = 4.69 105
2 = (22.75 536000
(1.987 708)) = 2.15 107
=((4.69 105) 0.00894 0.0673) ((2.15 107) 0.09584 0.06730.5)
0.008940.5
= 2.42 107
1 = 4129794
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52
Figure 18: Plot of 1/-RA v/s XA for third bed
FA0 = 36.355kmol/hr = 36.355/3600kmol/s = 0.0100986kmol/s
W = FA0 * Area under the curve= 0.0100986*181425
Weight of catalyst in first catalyst bed = 1832.14kg
13.2.4 Fourth Catalytic Bed
Table 19: Calculations of Fourth Catalytic Bed
T(K) XSO2 NSO2 NO2 NSO3 NN2 Total PSO2(atm) PO2(atm)
PSO3(atm)
700 0.946 22.27 261.63 395.22 3298.306 3977.442 0.0056 0.06578
0.099367
701 0.950 21.350 261.17 396.15 3298.306 3976.981 0.005369
0.065671 0.099611
702 0.953 19.851 260.42 397.64 3298.306 3976.231 0.004993
0.065495 0.100007
704 0.960 16.864 258.93 400.63 3298.306 3974.738 0.004243
0.065144 0.100796
705 0.963 15.376 258.18 402.12 3298.306 3973.994 0.003869
0.064969 0.101189
707 0.970 12.409 256.70 405.1 3298.306 3972.51 0.003124 0.06462
0.101974
708 0.97 10.931 255.964 406.57 3298.306 3971.771 0.002752
0.064446 0.102365
-
53
Sample Calculation:
For T = 704 K, T0 = 700 K, 2 = 0.96
704 =700 + 307.158 (2 0.96)
1 + 0.0248 (2 0.96)
2 = 0.96
2 = 417.501(1 0.96)
(1 0.05 0.96)
683
704
2 = 16.864
2 = 417.501 16.864 = 400.63
3 = 400.63
2 = 459.249 (400.63 2 ) = 258.93
= 16.864 + 258.93 + 400.63 + 3298.306 = 3974.738
2 = 2 =16.864
3974.738 = 0.004243
2 = 2 =258.93
3974.738 = 0.065144
3 = 3 =400.63
3974.738 = 0.100796
Table 20: Rate Calculations of Fourth Catalytic Bed
k1 k2 (-rA) (kmol/s-kg catalyst) 1/(-rA)
3.65E-05 1.39E-07 1.32E-07 7557000
3.76E-05 1.47E-07 1.3E-07 7713771
3.89E-05 1.55E-07 1.24E-07 8072529
4.14E-05 1.73E-07 1.07E-07 9315462
4.27E-05 1.83E-07 9.67E-08 10343043
4.55E-05 2.04E-07 6.97E-08 14342480
4.69E-05 2.15E-07 5.19E-08 19276120
-
54
Sample Calculation:
1 = (12.07 31000
(1.987 704)) = 4.14 105
2 = (22.75 536000
(1.987 704)) = 1.73 107
=((4.14 105)0.004243 0.065144) ((1.73 107) 0.100796
0.0651440.5)
0.0042430.5
= 1.07 107
1 = 9315462
Figure 19: Plot of 1/-RA v/s XA for fourth bed
FA0 = 22.27kmol/hr = 22.27/3600 kmol/s = 6.18611*10-3kmol/s
W = FA0 * Area under the curve = 6.18611*10-3*2474750
Weight of catalyst in first catalyst bed = 15309.078 kg
-
55
14. SUMMARY SHEET
1. Diameter of converter = 6 m.
2. Heat load for heat exchanger after first bed = -6361.36
kW
3. Heat load for heat exchanger after second bed = -231.106
kW
4. Heat load for heat exchanger after third bed = -554.68 kW
5. Design Pressure for reactor = 0.1114 N/mm2
6. Design Temperature for reactor =610C =883 K
7. Material of Construction =SS 304
14.1 First Catalytic Bed
Inlet temperature = 410C
Effluent temperature = 602C
Conversion = 0.66
Weight of catalyst = 29103.3 kg.
Heat duty required= -728.67 kW
14.2 Second Catalytic Bed
Inlet temperature = 438C
Effluent temperature = 498C
Conversion = 0.85
Weight of catalyst = 4950.878
Heat duty required= -192.1755 kW
14.3 Third Catalytic Bed
Inlet temperature = 432C
Effluent temperature = 444 C
Conversion = 0.96
Weight of catalyst = 1832.14 kg
Heat duty required= -7.7903 kW
-
56
14.4 Fourth Catalytic Bed
Inlet temperature = 427C
Effluent temperature = 435C
Conversion = 0.985
Weight of catalyst = 15309.078 kg
Heat duty required = -1.69 kW
-
57
15. CONCLUSION
The next step to the Contact Process is DCDA or Double Contact
Double Absorption. In this
process the product gases (SO2) and (SO3) are passed through
absorption towers twice to achieve
further absorption and conversion of SO2 to SO3 and production
of higher grade sulfuric acid.
SO2-rich gases enter the catalytic converter, usually a tower
with multiple catalyst beds, and are
converted to SO3, achieving the first stage of conversion. The
exit gases from this stage contain
both SO2 and SO3 which are passed through intermediate
absorption towers where sulfuric acid
is trickled down packed columns and SO3 reacts with water
increasing the sulfuric acid
concentration. Though SO2 too passes through the tower it is
unreactive and comes out of the
absorption tower.
This stream of gas containing SO2, after necessary cooling is
passed through the catalytic
converter bed column again achieving up to 99.8% conversion of
SO2 to SO3 and the gases are
again passed through the final absorption column thus resulting
not only achieving high
conversion efficiency for SO2 but also enabling production of
higher concentration of sulfuric
acid.
The industrial production of sulfuric acid involves proper
control of temperatures and flow rates
of the gases as both the conversion efficiency and absorption
are dependent on these.
-
58
16. REFERENCES
[1] Hill, Charles G. (1977). An Introduction to Engineering
Kinetics and Reactor Design, Third Edition,
506,507,509-513.
[2] Douglas K. Louie, (2005). Handbook of Sulphuric Acid
Manufacturing , Second Edition, 3-3, 3-8, 3-
10, 3-11, 3-14, 3-28, 3-78, 4-5 4-11, 4-58,6-16-3.
[3] Matt King, Michael Moats and William G.I. Davenport (2013).
Sulfuric Acid Manufacture (Second
Edition) Analysis, Control and Optimization, 11-14, 19-21,
73-80, 91-99, 229-234, 262-271.
[4] Levenspiel, Octave, (1999). Chemical Reaction Engineering,
Third Edition, 208, 209, 212, 394.
[5] Christie John Geankoplis, (2003). Transport Processes and
Separation Process Principles (Includes
Unit Operations) Fourth Edition, 22.
[6] Central Pollution Control Board Ministry of Environment and
Forests (May 2007). Comprehensive
Industry Document on Sulphuric Acid Plant, 3, 6-7, 10-14, 22,
25-28, 35-36, 38.
[7] Fogler, H.S., 1999. Elements of chemical reaction
engineering, Third edition. Prentice Hall, New
York.
[8] Bhatt, B.I., Vora, S.M., 1996. Stoichiometry. Third edition.
Mc-Graw Hill, New Delhi.
[9] Austin G.T., Shreve, 1984 Chemical process industries,
Fourth edition.Mc-Graw Hill, Singapore.
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59
17. APPENDIX (DERIVATIONS)
17.1 Derivation of relationship between equilibrium conversion
and temperature.
Reaction taking place in the reactor:
SO2 (g) + O2 (g) SO3 (g)
Rate Equation:
The kinetic data are represented by a rate expression of the
form
=122 2 32
0.5
20.5 (21)
that may be regarded as a degenerate form of typical
Hougen-Watson kinetics.
At equilibrium, the rate of forward reaction becomes equal to
the rate of backward reaction.
=122 2 32
0.5
20.5 = 0 (22)
Therefore at equilibrium,
1 22 = 2 320.5 (23)
The partial pressures of the various species are numerically
equal to their mole fractions since
the total pressure is one atmosphere. These mole fractions can
be expressed in terms of a single
reaction progress variable-the degree of conversion-as indicated
in the following mole table.[8]
V2O 5
-
60
Table 21: Table for mole fractions expressed in terms of
conversion
Component
Initial moles
(kmol/hr) Moles at conversion XA
Mole fraction at fractional
conversion XA
SO2 417.501 417.501 (1 ) 417.501 (1 )
4175.056 208.7505
O2 459.249 459.249 459.249
(417.501 )
2 208.7505
4175.056 208.7505
N2 3298.306 3298.306 3298.306
4175.056 208.7505
SO3 0.0 417.501 417.501
4175.056 208.7505
Total 4175.056 4175.056 208.7505 1.0
At equilibrium,
1 22 = 2 320.5 (24)
1 (417.501 (1 )
4175.056 208.7505 ) (
459.249 208.75054175.056 208.7505
) =
2 (417.501
4175.056 208.7505 ) (
459.249 208.75054175.056 208.7505
)0.5
(25)
-
61
=12
=(
417.501
4175.056208.7505 )
(417.501 (1)
4175.056208.7505 ) (
459.249208.7505
4175.056208.7505 )
0.5 (26)
=12
=
(1 )
20 2.2
(27)
17.2 Derivation of relationship between conversion and
temperature.
The differential form of energy balance equation for
one-dimensional, plug flow model, for
adiabatic operation is as follows, [9]
()
= () = () (28)
where,
= = superficial mass velocity, which does not vary along the
length of reactor.
= fluid density
= superficial velocity in axial direction.
= reaction rate expressed in pseudo homogeneous form (i.e.
number of moles transformed per
unit time per unit of total reactor volume)
= enthalpy change for the reaction at the indicated
conditions.
= bulk density of the catalyst (total mass of catalyst / total
volume of reactor)
=
= global reaction rate per unit mass of catalyst.
The equation for material balance is as follows:
()
= = (29)
where
= stoichiometric coefficient for reactant A (negative for
reactants) = 1 in this case
-
62
= (1 + )
(30)
= (1 )
(1 + )
(31)
Therefore
()
=
(32)
=
(33)
()
=
() (34)
Table 22: Mole percent of gases entering the converter
Component Inlet Moles
(Percent )
Molecular
Weight
SO2 10 64
O2 11 32
N2 79 28
Average molecular weight of the inlet gas:
Mavg = 0.10*64+ 0.11*32+0.79*28 = 32.04
For the temperature range of interest (875K to 683K), the heat
capacity per unit mass is
substantially independent of the conversion level. Hence, we
take the heat capacity as constant at
0.250 cal/gm-K.
( ) = ( ) (35)
Diameter of Converter is taken as 6m. [1]
-
63
() =
=4175.056 32.04
(62)
4
= 4731.1033
m2= 1.314
2
2 = =417.501 103
3600 (6)2
4
= 4.101688
2
Heat of reaction at temperature T = (H) = 24.60 + 1.99 x 10-3T
kcal/gmol
for T in degrees Kelvin.
1.314 0.250 ( ) = (24.60 1.99 x 103T) 4.101688( )
= + 307.158( )
1 + 0.0248( ) (36)