1. Department Of Chemical &
Environmental Engineering: 3rd
Year Individual Design Project
(Fixed Bed Catalytic Reactor)
AUTHOR: SAMUEL ESSIEN
APRIL 1, 2013
UNIVERSITY OF NOTTINGHAM
2. 1
Table of Contents
1. Executive Summary.........................................................................................................................2
2. Design Basis.....................................................................................................................................2
3. Design Constraints ..........................................................................................................................3
4. Environmental Considerations........................................................................................................3
5. Design Optimisation, Mechanical Design & Equipment Sizing .......................................................4
6. Estimated Cost Of Plant Section .....................................................................................................4
7. Design Alternatives & Alternate Methodologies............................................................................4
8. Start-up & Shutdown Procedure.....................................................................................................6
9. References ......................................................................................................................................7
PLANT SPECIFICATION SHEET..................................................................................................................8
i) Enthalpy Data................................................................................................................................14
ii) Determination of Thermodynamic Pathway & Bed Conversion Using Multivariable Matrix
Calculation ............................................................................................................................................15
iii) Bed Sizing Using Conversion Bed Conversion Data from Appendix 1ii) and Rate Equation
Supplied By Lucite.................................................................................................................................19
iv) Off-gas Cooling & Heating Requirements Between Beds (All Temperatures in Kelvin) ...............31
HEAT EXCHANGER 1 SPECIFICATION ....................................................................................................35
HEAT EXCHANGER 2 SPECIFICATION ....................................................................................................36
HEAT EXCHANGER 3 SPECIFICATION ....................................................................................................37
HEAT EXCHANGER 4 SPECIFICATION ....................................................................................................38
MATERIAL COSTING SHEET & OUTPUT INVENTORY.............................................................................40
BUILDING MATERIAL & ERECTION COSTS.............................................................................................41
START-UP PROCEDURE .........................................................................................................................42
SHUTDOWN PROCEDURE .....................................................................................................................42
3. 2
1. Executive Summary
In summary, the design that has been carried out has highlighted the potential socio-economic
benefits that this project could pose for the client. If this project is given commissioning, it will be
possible to reduce the annual SO2 off-gas emission by 99% saving millions of pounds in incurred fines
each year. Not only will this project save the stakeholders money (approximately over
£60,000,000.00 in savings on fines that will be incurred if no action is taken) but will also generate
annual revenues of over 30 million pounds (see Appendix 3 costing sheet). In order to meet this
target the design I carried out not only had to be numerically sound but also meet a high level of
accuracy. To do this, I made use of Microsoft Excel and set up a system of Vector operations to
model the thermodynamics of the reactions taking place (see Appendix 2). Because the oxidation
reaction of Sulphur dioxide is reversible, the dynamics of the backward and forward reactions had to
be modelled in order to find the overall conversion in each catalytic bed. In the end, it was possible
to convert 99.62% through optimisation of the catalytic reactor beds using the goal finder tool to
iterate between various operating parameters with the set design basis value specified in the task
hand-out document. Initially I considered various design layouts and various methods of calculation,
however as the design developed and I learned more and more about the underlying processes, I
was able to logically string together a feasible process including start-up, shut-down, emergency
protocol and process control procedures.
Fig 3.1 – Trade-off between fixed capital requirement of battery region and potential annual
earnings based on the NPV of inventory chemicals and saleable commodities produced during
catalytic conversion in the double contact process.
2. Design Basis
As specified in the design brief, my team has been given the task of designing a process to utilize
downstream flue gas containing a reasonable amount of sulphur dioxide and oxygen. In order to
utilize this gas in the most efficient and cost effective manner, we decided in Task 1 to make use of
the double contact process. The basis of the design and quantified outputs from each converter bed
has been summarised in the plant specification sheet. For simplicity each stream has been specified
with a number which is visible on the PNID diagram of my design battery region of the plant which
goes from the inlet region of Bed 1 all the way to the outlet of Bed 4. In addition to the Off-gas
produced in the upstream plant, we decided to oxidise an additional 412kgmol/h of Sulphur using a
sulphur burner to enrich the off-gas stream with an additional 499.4kgmol/h of SO2. The gas coming
£9,500,000.00
(Battery Region Fixed Capital)
£31,000,000.00
(Estimated Minimum Annual
Revenue)
4. 3
out of the burner and coming into the converter has a molar flow rate of 6025.57kgmol/h and an
initial concentration of %11.84 SO2 and 7.34% O2 (the rest of the stream is considered as non-
reacting components) it is also assumed that all the water in the stream has been taken out during
the gas drying stages upstream. The presence of water in the converter can lead to mist formation in
the bed, which can cause serious corrosion damage to the catalyst beds and internal fittings. For this
reason Stainless Steel was my choice of material to design the converter shell and process side
piping with. Carbon Steel is used for all of the service line piping and shell side material in the 4 heat
exchangers I have designed.
3. Design Constraints
The primary constraints in my design battery region based on my groups work in Task 1 and 2 were:
i) The Inlet Molar Flow which was specified at 6025.57kgmol/h from Tim Royston’s Burner
design.
ii) Absorption of a maximum of 98% SO3 in the interpass absorber leaving at least
13.57kgmol/h of SO3, 34.78kgmol/h of SO2 and 103.12kgmol/h of oxygen (in excess to
promote forward reaction)as feed gas to Bed 4.
iii) Constraints Set By Thermodynamics & Rate Equation
a) At steady state, each bed cannot achieve a conversion higher than its Equilibrium
Conversion; during steady state operation it is possible to achieve a maximum
conversion of 95% of the bed equilibrium conversion.
b) Excess Heat In Bed Promotes Backward Endothermic Reaction and Slows Down
Forward Exothermic Reaction (Le Chatelier’s Principle)
c) Each Bed cannot be operated above 900k as this causes thermal degradation of the
catalyst.
4. Environmental Considerations
In the converter, there is a likelihood of mist formation occurring if water vapour gets into the
packing bed. This mist is primarily sulphuric acid however trace amounts of sulphurous acid may
form also; the best way to deal with this risk is to have demister pads in each bed and a layer of fibre
brick to prevent corrosion. The environmental problem arises when the gas gets out of bed 4. If
there is mist present in the gas leaving bed 4, this may pose major environmental issues if it
manages to get out of the scrubber to the stack. The mist is toxic and highly corrosive and can
damage nearby forestry and contaminate nearby water bodies. Therefore it is imperative that the
inlet stream to the converter is dry off-gas.
The off-gas from bed 4 is sent to the secondary absorber (gas scrubber) and then sent through to the
gypsum tower where the remaining SO2 is reacted with lime water, to produce gypsum which is sold
to market. Limewater also reacts with the CO2 in the stream to form Calcium Carbonate. The
reaction for formation of gypsum from lime water is shown below:
Ca(OH)2(aq) + SO2(g) → CaSO3(s) + H2O(l)
Ca(OH)2 + CO2 → CaCO3 + H2O
5. 4
The primary benefit of gypsum tower is that it reduces the amount of SO2 and CO2 output to
atmosphere which will save the client money on fines incurred.
5. Design Optimisation, Mechanical Design & Equipment Sizing
The optimisation of all the equipment in my design has been carried out using the goal find tool,
information on the process used to optimise and size all of the equipment is shown in Appendix 2.
The dimensions obtained are shown in the plant specification sheet.
6. Estimated Cost of Plant Section
By carrying out a mechanical design of the converter I was able to determine the thickness and
therefore the weight of the shell. I was then able to find a cost estimate using the cost data in
Coulson & Richardson Volume 6. I carried out the same calculation for each of the equipment in the
battery region of my converter and was able to determine the fixed capital required to erect the all
of the equipment in my battery region. Appendix 3 shows the calculation spread sheet I put together
for costing. The PCE was found to be £3,956,418.12 and the total fixed capital required to build the
entire battery region is 9.5 million pounds.
7. Design Alternatives & Alternate Methodologies
I decided to make use of 25mm & 12mm Caesium Enriched daisy rings due to the efficiency they
offer even when dust and debris particles are present in the stream. They also offer the lowest
pressure drop per metre height of packing material. In order to prevent high levels of corrosion in
the converter, all of the fixed and welded fittings within the bed are made of high quality stainless
steel. Although carbon Steel with Brick Lining could be considered for low concentration off-gas
systems, the most efficient operation of the double contact process is achieved when stainless steel
is used as the material of choice for the converter shell and internal fittings within each bed. The
Table Below shows how Bed Height & Pressure Drop varies with Catalyst Packing Weight
Requirement for each type of packing.
Fig 1.1 – Variation of bed packing height and pressure loss for different packing types.
i) Nozzle Design Alternatives
The nozzle design can consist of different types of transitions used to reduce pressure losses as the
off-gas passes from the pipe out of the nozzle. There are 3 major types of transitions that can be
used in nozzles 1
; Asymmetrical Transition, Symmetric Transition and Gas Box Transition. These are
shown in the figure below:
1
Handbook of Sulphuric Acid Manufacturing Chapter 3 pg 36 -38
10mm Rashig
Ring
12mm
Daisy Ring 6mm Pellets
10mm Rashig
Ring
12mm Daisy
Ring 6mm Pellets
Bed 1 24.15 0.5651 0.6692 0.4986 0.0324 0.0320 0.0334
Bed 2 45.08 1.0546 1.2489 0.9305 0.0458 0.0452 0.0472
Bed 3 52.20 1.2212 1.4462 1.0775 0.0532 0.0525 0.0548
Bed 4 32.43 0.7587 0.8985 0.6695 0.0413 0.0408 0.0425
Weight of
Catalyst / mT
Bed Height Requirement / m Pressure Loss / Kpa
6. 5
Tw
Fig 1.2- (Left to Right) Transition Design Alternatives – Symmetrical Transition, Asymmetrical Transition & Gas Box Transition.
The symmetrical transition is the easiest to design and only require stiffening ribs to withstand the
pressure of the incoming gas. They also produce the lowest head loss in the gas and they have
minimal bends. The Gas box transition is used when there is a duct immediately after the converter
inlet or outlet, the shell of the converter must be designed to handle the stress in the walls near by
the gas box transition. There is also a higher pressure loss due to the bend present in the gas box
design, I decided to go for a simple symmetrical design that is easy to cost and requires little
reinforcement to the shell structure other than welds are around the ribbing to the shell and rim of
the nozzle.
ii) Thermal Insulation Alternatives
In my design I decided to place fibre brick on the inside of the beds to reduce corrosion damage over
the long haul of the plant lifecycle and to provide additional thermal insulation as a safety feature of
the design. I also decided to place mineral wool insulation on the outside to ensure that the outside
of the converter is not so hot that it causes serious burns to workers if touched. The wall thickness is
specified in the plant specification sheet.
An alternative would be to place mineral wool on both sides, as it has a high thermal stability up to
1200o
C but this would be rather expensive to design and build because the mineral wool would
require fastening to the shell at numerous points and offers no mechanical support to the structure.
The brick however is inexpensive and can offer minor structural support to the internal fittings and
beds. For this reason I chose to use fibre brick.
The diagram shows the layers of the walls inside the converter.
Ti
iii) Gas Distribution Design Alternatives
Ts Ts
i Ti
25o
C
Fibre Brick
Stainless Steel
Mineral Wool
Fig 1.3 – Showing the Thermal Insulation Used in The Converter.
Fibre brick provided additional structural integrity to the bed and
protects the stainless steel shell against corrosion.
7. 6
In my design, gas comes into the central pipeline present in the central core and is then
distributed radially through 4 pipes of smaller diameter each fitted with a gas nozzle to
efficiently disperse gas through each bed. The flow path of the gas is shown below:
Fig 1.4 – Showing the expected radial flow gas distribution in the catalytic
converter beds. This design allows more efficient conversion within the bed
because it prevents thermal gradients from developing within the bed which
reduce conversion efficiency and can lead to structural damage within the
converter over time. Radial flow also increases the gas turbulence in the bed
which is essential for efficient bed operation.
A key safety feature of this design is that the entire
column can be purged of Off-gas through the central
pipeline in the event of an emergency or for routine
maintenance. Individual beds can also be bypassed if
required; for example if the client decided to carry out
single contact acid making, bed 4 can be shut-down by shutting off the gas supply valve to bed 4.
This can allow the client to utilise both single and double contact acid making during as required
during peak and off-peak periods of market demand to better. Alternative methods of gas
distribution are shown in the figure below:
8. Start-up & Shutdown Procedure
i) Start-up Procedure (See Appendix 4)
In order for the beds to operate efficiently, the catalyst bed must be heated to its active
temperature at which point the caesium-vanadium oxide layer is in a molten state (roughly 593K-
653K depending on the catalyst product used in the bed) . In the case of this design, hot air is blown
into each catalyst bed from the burner line. This air must be sufficiently hot to heat the bed to its
steady state operating temperature. The air must also be dry and contain no water vapour to avoid
wetting the packing which leads to acid mist formation in the beds during steady state operation.
The air is initially heated by hot fossil fuel combustion (using a fuel fired gas heat exchanger), which
Fig 1.4a) Gas Nozzle Distribution
Fig 1.4b) Baffle Plate Gas Distribution
Fig 1.4c) Diffuser Gas Distribution
10. 9
Choice Of Material & Design Alternatives –
Appendix 1
Fig 1.1 – Catalyst Packing Alternatives courtesy BASF & Haldor Topsoe
In order to build an efficient converter an ignition layer of VK69 is used on the upper layers of each
catalyst bed. The purpose of this layer is to generate heat in the packing to ignite the layers of lower
activity packing below . The thickness of this layer has been estimated to be 20 per cent of packing
height within each bed, this is because in order to find the actual thickness a dynamic model would
have to be set up using a CFD program or mathematical Fourier Heat model to model how the heat
propagates through the packing. I decided to make use of 25mm & 12mm daisy rings due to the
efficiency they offer even when dust and debris particles are present in the stream. They also offer
the lowest pressure drop per metre height of packing material.
In order to prevent high levels of corrosion in the converter, all of the fixed and welded fittings
within the bed are made of high quality stainless steel. Although carbon Steel with Brick Lining could
be considered for low concentration off-gas systems, the most efficient operation of the double
contact process is achieved when stainless steel is used as the material of choice for the converter
shell and internal fittings within each bed.
iv) Nozzle Design Alternatives
The nozzle design can consist of different types of transitions used to reduce pressure losses as the
off-gas passes from the pipe out of the nozzle. There are 3 major types of transitions that can be
used in nozzles 2
; Asymmetrical Transition, Symmetric Transition and Gas Box Transition. These are
shown in the figure below:
2
Handbook of Sulphuric Acid Manufacturing Chapter 3 pg 36 -38
11. 10
Tw
Fig 1.2- (Left to Right) Transition Design Alternatives – Symmetrical Transition, Asymmetrical Transition & Gas Box Transition.
The symmetrical transition is the easiest to design and only require stiffening ribs to withstand the
pressure of the incoming gas. They also produce the lowest head loss in the gas and they has
minimal bends. The Gas box transition is used when there is a duct immediately after the converter
inlet or outlet, the shell of the converter must be designed to handle the stress in the walls near by
the gas box transition. There is also a higher pressure loss due to the bend present in the gas box
design, I decided to go for a simple symmetrical design that is easy to cost and requires little
reinforcement to the shell structure other than welds are around the ribbing to the shell and rim of
the nozzle.
v) Thermal Insulation Alternatives
In my design I decided to place fibre brick on the inside of the beds to reduce corrosion damage over
the long haul of the plant lifecycle and to provide additional thermal insulation as a safety feature of
the design. I also decided to place mineral wool insulation on the outside to ensure that the outside
of the converter is not so hot that it causes serious burns to workers if touched.
An alternative would be to place mineral wool on both sides, as it has a high thermal stability up to
1200o
C but this would be rather expensive to design and implement as the mineral wool would
require fastening to the shell at numerous points and offers no mechanical support to the structure.
The brick however is inexpensive and can offer minor structural support to the internal fittings and
beds. For this reason I chose to use fibre brick.
The diagram shows the layers of the walls inside the converter.
Ti
Ts Ts
i Ti
25o
C
Fibre Brick
Stainless Steel
Mineral Wool
12. 11
vi) Gas Distribution Design Alternatives
In my design, gas comes into the central pipeline present in the central core and is then
distributed radially through 4 pipes of smaller diameter each fitted with a gas nozzle to
efficiently disperse gas through each bed. The flow path of the gas is shown below:
Fig 1.3 – Showing the expected radial flow gas distribution in the catalytic converter beds. This design allows more efficient conversion
within the bed because it prevents thermal gradients from developing within the bed which reduce conversion efficiency and can lead
to structural damage within the converter over time.
A key safety feature of this design is that the entire can be purged of Off-gas through the central
pipeline in the event of an emergency or for routine maintenance. Individual beds can also be
bypassed if required; for example if the client decided to carry out single contact acid making, bed 4
can be shut-down by shutting off the gas supply valve to bed 4. This can allow the client to utilise
both single and double contact acid making during as required during peak and off-peak periods of
market demand to better. Alternative methods of gas distribution are shown in the figure below:
Fig 1.4a) Gas Nozzle Distribution
Fig 1.4b) Baffle Plate Gas Distribution
Fig 1.4c) Diffuser Gas Distribution
13. 12
Simultaneous Energy & Mass Balance Calculations Using the Solver
Tool in Microsoft Excel - APPENDIX 2
In order to set up a simultaneous equation of each catalyst bed, a set of simultaneous equations
have to be set up.
For the above set of equations (set up as a 7-7 Matrix (A) highlighted in BLUE and a 7-1 (B) Matrix
highlighted in RED) and the numerical solution Matrix (C) the following set of mathematical
equations apply. C is a single column matrix which when multiplied by A will give the single column
matrix B.
∴ 𝐴 × 𝐶 = 𝐵
∴ 𝐶 = 𝐴−1
× 𝐵 ------ (1)
Where A-1
is the inverse matrix of A.
Hence for a simplified set of arrays (5-5 and 5-1’s say) the following applies:
For a simple reaction involving monoatomic elements of the form J+K=I, where I=JK. If the bottom
row of matrix B is the enthalpy of each individual component at a specified inlet & outlet
temperature then A, B and C become:
Let 𝐶 =
[
𝑥𝑗𝑖
𝑥 𝑘𝑖
𝑥𝑖𝑜
𝑥𝑗𝑜
𝑥 𝑘𝑜]
where x denotes the mole fraction of each component specified.
Let 𝐵 =
[
∑ 𝑟1
∑ 𝑟2
∑ 𝑟3
∑ 𝑟4
∑ ∆ℎ]
where n denotes a numerical term which is the total of sum of a given row in Matrix
A (and not A-1
)
Bed 1
Numerical Term SO2 in O2 in CO2 in N2 in SO3 out SO2 out O2 out CO2 out N2 out
Feed SO2 kg-mole 0.121382989 1 0 0 0 0 0 0 0 0
Feed O2 kg-mole 0.060749883 0 1 0 0 0 0 0 0 0
Feed CO2 kg-mole 0.02130209 0 0 1 0 0 0 0 0 0
Feed N2 kg-mole 0.78817734 0 0 0 1 0 0 0 0 0
S Balance 0 -1 0 0 0 1 1 0 0 0
O Balance 0 -2 -2 -2 0 3 2 2 2 0
C Balance 0 0 0 -1 0 0 0 0 1 0
N Balance 0 0 0 0 -2 0 0 0 0 2
Enthalpy Balance
per kg-mole 0 280.7535 -10.868 378.2335 -10.418 -356.87887 -268.166 18.92988 -366.037 17.94274
Feed Gas Temp 650
Product Temp 891.953056
14. 13
𝐴 =
[
1 0 0 0 0
0 1 0 0 0
−1
0
ℎ𝑗𝑖
0
−1
ℎ 𝑘𝑖
1 1 0
1 0 1
ℎ𝑖𝑜 ℎ𝑗𝑜 ℎ 𝑘𝑜]
Equation (1) then becomes:
𝑪 =
[
1 0 0 0 0
0 1 0 0 0
−1
0
ℎ𝑗𝑖
0
−1
ℎ 𝑘𝑖
1 1 0
1 0 1
ℎ𝑖𝑜 ℎ𝑗𝑜 ℎ 𝑘𝑜]
−1
×
[
∑ 𝑟1
∑ 𝑟2
∑ 𝑟3
∑ 𝑟4
∑ ∆ℎ]
∴ 𝑪 =
𝟏
𝑫𝒆𝒕(𝑨)
× 𝑨 𝑻
×
[
∑ 𝑟1
∑ 𝑟2
∑ 𝑟
3
∑ 𝑟
4
∑ ∆ℎ]
𝒘𝒉𝒆𝒓𝒆 𝑨 𝑻
𝒊𝒔 𝒕𝒉𝒆 𝒕𝒓𝒂𝒏𝒔𝒑𝒐𝒔𝒆 𝒐𝒇 𝒕𝒉𝒆 𝒎𝒂𝒕𝒓𝒊𝒙 𝑨
It should be noted that the enthalpy at any temperature has been found using formation data, since
hT=Δfh298K + (hT - Δfh298K) provided the temperatures are within a one phase region of the component,
a linear relationship for hT against T can be determined. This allows the energy balance to be carried
out on the system by multiplying the input & output temperature by each hT to determine the
enthalpy of each component. Hence at any temperature the %conversion of any reactant can be
plotted against the outlet temperature of the stream this is called the enthalpy pathway.
∴ 𝐹𝑜𝑟 𝐸𝑥𝑎𝑚𝑝𝑙𝑒 → ℎ𝑗𝑖
= ℎ 𝑇𝑗 (𝑇𝑖) 𝒘𝒉𝒆𝒓𝒆 𝒉 𝑻𝒋 𝒊𝒔 𝒕𝒉𝒆 𝒆𝒏𝒕𝒉𝒂𝒍𝒑𝒚 𝒐𝒇 𝒄𝒐𝒎𝒑𝒐𝒏𝒆𝒏𝒕 𝒋 𝒂𝒕 𝒕𝒉𝒆 𝒊𝒏𝒍𝒆𝒕 𝑻𝒆𝒎𝒑𝒆𝒓𝒂𝒕𝒖𝒓𝒆 𝑻𝒊
The numerical solution of the matrix C can therefore be found for any reaction of n-components
provided that the total degree of freedom of the system is zero. That is in other words, the total
number of species involved (either reacting or being formed) is equal to the total number of
simultaneous equations that can be set up to solve the problem. Using Microsoft Excel, the
equations above have been used to determine the operational conditions of the Packed Bed Reactor
along with the exit gas composition from each bed.
The command MMULT(MINVERSE(…. : ….), …. : ….) then CTRL+SHIFT+ENTER with a column of row
length C, produces the solution of matrix C required. The same set of simultaneous can be used to
carry out a similar hand calculation using an Augmented Matrix of [A|B] and carrying Gaussian
Elimination to transform [A|B] into [IA
|C] where IA
is an identity matrix of the same dimensions as
A. The goal finder tool can then be used to determine optimal parameters of the Bed Reactor.
J in Feed (Row 1), K in Feed (Row 2)
Stoichiometric Number of Atoms of J (row 3)
and K (row 4) in components J,K and I
component enthalpies (Row 5)
15. 14
Temperature, K
298.15 0
600 8.894 8.894
700 11.937 11.937
800 15.046 15.046
900 18.223 18.223
0.031096 -9.797 H = 0.03110*T - 9.797 (Enthalpy at Temp T)
Temperature, K
298.15 -393.522
600 12.907 -380.615
700 17.754 -375.768
800 22.806 -370.716
900 28.03 -365.492
0.050421 -410.9635 H = 0.05041*T - 411.0 (Enthalpy at Temp T)
Temperature, K
298.15 -813.989
300 0.257 -813.732
400 15.112 -798.877
0.14855 -858.297 H = 0.1486*T - 858.3 (Enthalpy at Temp T)
H₂SO₄(l)
H83PDX Enthalpy Data (Units in all enthalpy colums)
N₂ (g)
CO₂ (g)
1 H83PDX Enthalpy Data (Units in all enthalpy colums)
2
3 Temperature, K
4 298.15 -395.765
5 600 18.107 -377.658
6 700 24.997 -370.768
7 800 32.160 -363.605
8 900 39.531 -356.234
9 0.07144 -420.6 H = 0.07144*T - 420.6 (Enthalpy at Temp T)
10
11
12 Temperature, K
13 298.15 -296.842
14 600 13.544 -283.298
15 700 18.548 -278.294
16 800 23.721 -273.121
17 900 29.023 -267.819
18 0.05161 -314.3 H = 0.05161*T - 314.3 (Enthalpy at Temp T)
19
20 O₂ (g)
21 Temperature, K
22 298.15 0
23 600 9.244 9.244
24 700 12.499 12.499
25 800 15.835 15.835
26 900 19.241 19.241
27 0.033327 -10.7905 H = 0.03332*T - 10.79 (Enthalpy at Temp T)
SO₃ (g)
SO₂ (g)
H83PDX Enthalpy Data (Units in all enthalpy colums)
Temperature, K
298.15 -285.83
300 0.139 -285.691
320 1.646 -284.184
340 3.153 -282.677
360 4.664 -281.166
380 6.182 -279.648
400 7.711 -278.119
0.075684286 -308.4036667 H = 0.07568*T - 308.4 (Enthalpy at Temp T)
H83PDX Enthalpy Data (Units in all enthalpy colums)
Temperature, K
298.15 0 0
300 27700.99743 0.86376668
400 65408.7142 2.03956078
500 127527.5526 3.976537344
600 220160.934 6.865011973
700 349412.2181 10.89529835
800 521384.7095 16.25770844
900 742181.6638 23.14255266
1000 1017906.293 31.74014011
1100 1354661.774 42.24077872
1200 1758551.25 54.83477549
1300 2235677.843 69.71243663
1400 2792144.654 87.06406779
1500 3434054.774 107.0799742
1600 4167511.287 129.9504611
1700 4998617.276 155.8658334
1800 5933475.832 185.0163964
1900 6978190.058 217.5924558
2000 8138863.074 253.7843179
2100 9421598.028 293.7822896
2200 10832498.1 337.776679
2300 12377666.49 385.9577952
2400 14063206.47 438.5159487
2500 15895221.35 495.6414515
0.203393799 -135.8052485 H = 0.203393799*T - 135.805 (Enthalpy at Temp T)
H₂O (l)
S(l)
The use of the matrix function and goal seek tool on excel allows the numerical solution of the bed
conversion to be obtained to an extremely high accuracy. The goal finder tool numerically iterates a
required equation in order to find numerical solutions for a set of required independent variables in
a given problem. This allows for processes such as plant optimisation to be carried out with ease to a
high degree of accuracy. The following figures show how I utilised matrix multiplication and the goal
seek tool to produce my results.
i) Enthalpy Data
In order to produce a model on excel the enthalpy-temperature relationship of the Off-gas had to be
determined, in order to achieve this I collated the enthalpy of each components at a range of values
and then linearized the results. The enthalpy data of each component that is in the OFF-GAS is
shown below, all units of enthalpy are in MJ/kg-mol :
Table 1.1-Linearised Enthalpy Temperature
Relationship. It can be assumed that within a
single phase temperature region, the enthalpy of
each off-gas component has an approximately
linear variation with temperature3
.
3
Sulphuric Acid Manufacture, W. Davenport and M. King – pg. 318-320 (table G1).
16. 15
0
10
20
30
40
50
60
70
80
90
100
600 700 800 900 1000
%Conversion
Temperature K
Plot of θᵗ Against Tc (Equillibrium
Curve)
θᵗ
Thermodynamic
Pathway
Equillibrium
Conversion
Curve
Operating
Temperature
Limit 900K
ii) Determination of Thermodynamic Pathway & Bed Conversion
Using Multivariable Matrix Calculation
As discussed at the start of appendix one, excel allows for simultaneous equations to be solved using
matrix operators in the function window. The solution of these calculations is a single vector column
matrix that contains the inlet and outlet molar flow rates of each component present in them. The
resulting molar flow of SO2 in and out of the bed is then used to find the bed conversion achieved at
a specified outlet and inlet temperature, this is known as the thermodynamic pathway. The
thermodynamic pathway shows how the forward reaction progresses with temperature for a
specified inlet gas Off-gas composition, it only takes into account the forward reaction that occurs in
the catalyst bed. The conversion of SO2 into SO3 is a reversible reaction and because the reaction is
exothermic, heat is released as SO2 is oxidised to SO3. This heat must be removed to maintain the
reactor at a steady operating temperature. The catalytic reaction requires a Vanadium (V) Oxide or
Titanium based catalyst and is a fast reaction occurring instantaneously. Hence the gas residence
time within the reactor is generally designed to be no more than 2-4 seconds. The equation of the
forward oxidation reaction is shown below:
2𝑆𝑂2(𝑔) + 𝑂2(𝑔)
700−900𝑘 𝑉2 𝑂5 𝑏𝑎𝑠𝑒𝑑 𝐶𝑎𝑡𝑎𝑙𝑦𝑠𝑡
→ 2𝑆𝑂3(𝑔) 𝛥𝐻 = −100𝑀𝐽 𝑝𝑒𝑟 𝐾𝑔𝑚𝑜𝑙 𝑆𝑂2
As discussed earlier, heat generated from the oxidation reaction must be removed to maintain each
bed at a steady state temperature. The catalytic bed is assumed to operate under adiabatic
conditions. However, because the dynamic reaction is reversible each bed can only reach a
maximum equilibrium conversion along a thermal pathway; temperature increase above this point
will favour the backward reaction and result in the thermal decomposition of SO3 back into SO2 and
oxygen. This is shown below:
17. 16
Above 900K the catalyst begins to thermally decompose, reducing the efficiency of the bed limiting
the conversion of the reaction further. Hence as stated previously the maximum attainable
conversion in the bed is the equilibrium conversion. The equations below are used to determine the
equilibrium conversion at any temperature based on the inlet SO2 and O2 concentration assuming all
other components are inert.
Fig 1.1 – Spread sheet showing simultaneous mass and energy balance calculation carried out using Microsoft Excel matrix operator
commands and the goal seek tool to determine the optimal conversion in bed 1 at the specified inlet temperature of 653K
The goal finder is set to find the intercept of the thermodynamic (heat-up) pathway of the reaction
and the equillibrium curve. Since the maximum attainable composition is the equillibrum
18. 17
Equillibrium Gas Composition
e, Volume% SO₂ = 11.84307023 f, volume %O₂ = 7.34428386
% Volume
Remainder
(inerts) = 80.812646
Pₑ = Equillibrium Pressure (Total System Pressure), bar = 1.2
θˢ = Suggested Equillibrium Curve Intercept (Bed Conversion)
θᵗ=Equillibrium Conversion 64.070448
64.07044839
Suggested Catalyst Bed Operating Temperature 872.02135
872.021352
Feed Temperature 653
Bed 1 Numerical Term SO2 in O2 in CO2 in N2 in SO3 out SO2 out O2 out CO2 out N2 out
Feed SO2 kg-mole 713.6127531 1 0 0 0 0 0 0 0 0
Feed O2 kg-mole 442.5351301 0 1 0 0 0 0 0 0 0
Feed CO2 kg-mole 90.21411332 0 0 1 0 0 0 0 0 0
Feed N2 kg-mole 4779.210253 0 0 0 1 0 0 0 0 0
S Balance 0 -1 0 0 0 1 1 0 0 0
O Balance 0 -2 -2 -2 0 3 2 2 2 0
C Balance 0 0 0 -1 0 0 0 0 1 0
N Balance 0 0 0 0 -2 0 0 0 0 2
Enthalpy Balance
per kg-mole 0 280.59867 -10.96796 378.08227 -10.5113 -358.30279 -269.195 18.26575 -367.041 17.32286
Feed Gas Temp 653
SO2 in 713.6127531 Product Temp 872.021352
O2 in 442.5351301
N2 in 90.21411332
CO2 in 4779.210253
SO3 out 457.2167273
SO2 out 256.3960259
O2 out 213.9267664
CO2 out 90.21411332
N2 out 4779.210253
Goal Finder 0.000
Catalyst Bed 1 Composition
Overall %
Conversion 64.0707058
Bed 1 %
Conversion 64.0707058
Specified Feed Gas Composition
e, Volume% SO₂ = 11.84307023
f,
volume
%O₂ = 7.344284
% Volume
Remainder
(inerts) = 80.81265
Pₑ = Equillibrium Pressure (Total System Pressure), bar = 1.2
1.2
θˢ = Suggested Equillibrium Curve Intercept (Bed Conversion)
θᵗ=Equillibrium Conversion 89.15023
89.15022908
Suggested Catalyst Bed Operating Temperature 770.3803
770.3802795
Feed Temperature 688
Bed 2 Numerical term SO3 in SO2 in O2 in CO2 in N2 in SO3 out SO2 out O2 out CO2 out N2 out
Feed SO3 kg-mole 457.2167273 1 0 0 0 0 0 0 0 0 0
Feed SO2 kg-mole 256.3960259 0 1 0 0 0 0 0 0 0 0
Feed O2 kg-mole 213.9267664 0 0 1 0 0 0 0 0 0 0
Feed CO2 kg-mole 90.21411332 0 0 0 1 0 0 0 0 0 0
Feed N2 kg-mole 4779.210253 0 0 0 0 1 0 0 0 0 0
S Balance 0 -1 -1 0 0 0 1 1 0 0 0
O Balance 0 -3 -2 -2 -2 0 3 2 2 2 0
C Balance 0 0 0 0 -1 0 0 0 0 1 0
N Balance 0 0 0 0 0 -2 0 0 0 0 2
Enthalpy Balance
per kg-mole 0 371.4493 278.7923 -12.13416 376.3179 -11.5998 -365.564 -274.441 14.87907 -372.165 14.16183
Catalyst Bed 2 Composition Inlet Temp 688
SO3 in 457.2167273 Outlet Temp 770.3802795
SO2 in 256.3960259
O2 in 213.9267664
CO2 in 90.21411332
N2 in 4779.210253
SO3 out 629.8262346
SO2 out 83.78651854
O2 out 127.6220128
CO2 out 90.21411332
N2 out 4779.210253
Goal Finder 0.000
88.25883
Overall %
Conversion
67.32144
Bed 2 %
Conversion
composition of the flue gas mixture at a given temperature. This is all done on microsoft excel and
the following spreadsheets were obtained:
Fig 1.1 – Spread sheet showing simultaneous mass and energy balance calculation carried out using Microsoft Excel matrix operator
commands and the goal seek tool to determine the optimal conversion in bed 1 at the specified inlet temperature of 653K
Fig 1.2 – Spread sheet showing simultaneous mass and energy balance calculation carried out using Microsoft Excel matrix operator
commands and the goal seek tool to determine the optimal conversion in bed 2 at the specified inlet temperature of 688K
19. 18
Fig 1.3 – Spread sheet showing simultaneous mass and energy balance calculation carried out using Microsoft Excel matrix operator
commands and the goal seek tool to determine the optimal conversion in bed 3 at the specified inlet temperature of 700K
The outlet gas from bed 3 is cooled to 480K and sent to the Interpass Absorber, with a correctly
optimised design it is possible to absorb 98% of the SO3 present in this stream and return off-gas at
350K to the catalytic reactor which must then be heated up to 683K before it is fed into the 4th
reactor bed. The inlet off-gas has a different equilibrium composition as a result of the change in
partial pressure of the components in the off-gas that occurs during SO3 absorption in the interpass
absorber. This allows a high conversion of 92.12% to be achieved in bed 4.
It should be noted that a fraction of SO2 and CO2 in the off-gas passing through the Interpass
Absorber will be absorbed during gas absorption because it has a similar solubility as SO3. This is
trace amounts however and therefore have been omitted within my design calculations.
Also the presence of CO2 slightly reduces the conversion achieved in each catalytic converter by
roughly 1 per cent. However for ease of calculation again this can be omitted. This is because the
specific heat capacity of CO2 is much larger than N2 therefore the presence of CO2 reduces the
sensible heat supplied by the inlet gas that is available for the reaction. The higher the
concentration of CO2 in the gas, the lower the thermal efficiency off each Bed.
Specified Feed Gas Composition
e, Volume% SO₂ = 0.34 f, volume %O₂ = 2.8
% Volume
Remainder
(inerts) = 96.86
Pₑ = Equillibrium Pressure (Total System Pressure), bar = 1.2
1.2
θˢ = Suggested Equillibrium Curve Intercept (Bed Conversion)
θᵗ=Equillibrium Conversion 96.08665136
96.08665136
Suggested Catalyst Bed Operating Temperature 723.3557332
723.3557332
Feed Temperature 700
Bed 3 Numerical term SO3 in SO2 in O2 in CO2 in N2 in SO3 out SO2 out O2 out CO2 out N2 out
Feed SO3 kg-mole 629.8262346 1 0 0 0 0 0 0 0 0 0
Feed SO2 kg-mole 83.78651854 0 1 0 0 0 0 0 0 0 0
Feed O2 kg-mole 127.6220128 0 0 1 0 0 0 0 0 0 0
Feed CO2 kg-mole 90.21411332 0 0 0 1 0 0 0 0 0 0
Feed N2 kg-mole 4779.210253 0 0 0 0 1 0 0 0 0 0
S Balance 0 -1 -1 0 0 0 1 1 0 0 0
O Balance 0 -3 -2 -2 -2 0 3 2 2 2 0
C Balance 0 0 0 0 -1 0 0 0 0 1 0
N Balance 0 0 0 0 0 -2 0 0 0 0 2
Enthalpy Balance
per kg-mole 0 370.592 278.173 -12.534 375.713 -11.973 -368.923 -276.868 13.31221 -374.536 12.69936
Catalyst Bed 2 Composition Feed Gas Temp 700
SO3 in 629.8262346 Product Temp 723.3557332
SO2 in 83.78651854
O2 in 127.6220128
N2 in 90.21411332
CO2 in 4779.210253
SO3 out 678.8297688
SO2 out 34.78298436
O2 out 103.1202457
CO2 out 90.21411332
N2 out 4779.210253
Goal Finder 0.000
Bed 3 %
Conversion 58.48618016
Overall %
Conversion 95.12578997
20. 19
Fig 1.3 – Spread sheet showing simultaneous mass and energy balance calculation carried out using Microsoft Excel matrix operator
commands and the goal seek tool to determine the optimal conversion in bed 4 at the specified inlet temperature of 683K
iii) Bed Sizing Using Conversion Bed Conversion Data from
Appendix 1ii) and Rate Equation Supplied By Lucite
The following equations were supplied by Lucite for bed sizing:
Reaction rate R (kmol/h per kg catalyst)
f
PkPk
kPP
P
PPk
R
SOSO
eqOSO
SO
SOO
2
32
5.01
32
22
3
22
1
1
Specified Feed Gas Composition
e, Volume% SO₂ = 0.692763356 f, volume %O₂ = 2.053818
% Volume
Remainder
(inerts) = 97.25341841
Pₑ = Equillibrium Pressure (Total System Pressure), bar = 1.2
1.2
θˢ = Suggested Equillibrium Curve Intercept (Bed Conversion)
θᵗ=Equillibrium Conversion 96.9683207
96.9683207
Suggested Catalyst Bed Operating Temperature 702.8362116
702.8362116
Feed Temperature 683
Bed 4 Numerical term SO3 in SO2 in O2 in CO2 in N2 in SO3 out SO2 out O2 out CO2 out N2 out
Feed SO3 kg-mole 13.57659538 1 0 0 0 0 0 0 0 0 0
Feed SO2 kg-mole 34.78298436 0 1 0 0 0 0 0 0 0 0
Feed O2 kg-mole 103.1202457 0 0 1 0 0 0 0 0 0 0
Feed CO2 kg-mole 90.21411332 0 0 0 1 0 0 0 0 0 0
Feed N2 kg-mole 4779.210253 0 0 0 0 1 0 0 0 0 0
S Balance 0 -1 -1 0 0 0 1 1 0 0 0
O Balance 0 -3 -2 -2 -2 0 3 2 2 2 0
C Balance 0 0 0 0 -1 0 0 0 0 1 0
N Balance 0 0 0 0 0 -2 0 0 0 0 2
Enthalpy Balance
per kg-mole 0 371.80648 279.0504 -11.96756 376.56997 -11.4443 -370.389 -277.927 12.6285 -375.57 12.06121
Catalyst Bed 2 Composition Feed Gas Temp 683
SO3 in 13.57659538 Product Temp 702.8362116
SO2 in 34.78298436
O2 in 103.1202457
N2 in 90.21411332
CO2 in 4779.210253
SO3 out 45.61863211
SO2 out 2.740947623
O2 out 87.09922733
CO2 out 90.21411332
N2 out 4779.210253
Goal Finder 0
Overall %
Conversion 99.61590546
Bed 4 %
Conversion 92.11986069
21. 20
T
ek
5473
16.12
1
T
ek
52596
45.71
3
T
ek
8028
953.9
2
7379.4
5006
10 T
eqk
f is the effectiveness factor of the catalytic bed, and is typically ~0.6.
Because the conversion achieved within the bed is a function of the inlet and outlet temperature ie
X=f(T), using the equations above it was possible to put together the stoichiometric table below:
𝐴 +
1
2
𝐵
𝑅𝑒𝑣𝑒𝑟𝑠𝑖𝑏𝑙𝑒
⇔ 𝐶
In Change Remainder Partial Pressure
NAO −𝑁𝐴𝑂 𝑋 𝑁𝐴𝑂(1 − 𝑋) 𝑃𝐴𝑂(1 − 𝑋)
NBO
−𝑁𝐴𝑂
1
2
𝑋 𝑁𝐴𝑂(𝜃 𝐵𝑂 −
1
2
𝑋) 𝑃𝐴𝑂(𝜃 𝐵𝑂 −
1
2
𝑋)
NCO +𝑁𝐴𝑂 𝑋 𝑁𝐴𝑂(𝜃 𝐵𝑂 + 𝑋) 𝑃𝐴𝑂(𝜃 𝐵𝑂 + 𝑋)
NIO ---------------------------- 𝑁𝐴𝑂 𝜃𝐼𝑂 𝑃𝐴𝑂 𝜃𝐼𝑂
The conversion is determined from the thermodynamic pathway therefore each partial pressure is
now a function of temperature also therefore using the specified values of X determined by Toutlet the
pressure drop parameter, weight of catalyst and height can be found for a specified bed diameter.
The typical height of packing is shown below:
Fig 1.4 – Typical Bed Thickness and Converter Diameters of a Converter. The Optimal Diameter is chosen using the goal finder to iterate
for the lowest pressure drop across each bed height. The diameter of the converter I designed has a diameter of 12.008m and an inner
core of diameter 5m.
22. 21
Ergun Equation Parameters
D G Dp gas ρ gas μ
6 0.6996815 0.025 0.529671 0.000055
Bed 1Sizing
T X(T) % K1= K2= K3= Keq= Pa(X(T)) Pb(X(T)) Pc(X(T)) R(T) 1/-R(T)
Area Under
Curve
Catalyst Weight
/Metric Tonnes
Catalyst Bed
Height m
Pressure
Drop Kpa
653 0 43.758374 10.39190944 8912.197 847.7274 0.142117 0.088131 0 -0.0536 18.6559 0.000000000 0.00 0.00 0.0000
660 0.0203 47.826303 9.12151915 3792.987 702.9889 0.139231 0.086688 0.002886 -0.00198 504.3792 5.310636868 3.79 0.10 0.0050
665 0.034817 50.90315 8.32431251 2083.343 616.4808 0.137169 0.085657 0.004948 -0.00231 432.0529 12.10445179 8.64 0.24 0.0114
670 0.0493311 54.127549 7.607157286 1154.581 541.6787 0.135106 0.084626 0.007011 -0.00362 276.0271 17.24302405 12.30 0.34 0.0163
680 0.1074 61.036411 6.378177211 363.9654 420.5956 0.126848 0.080497 0.015268 -0.00688 145.2533 29.48225225 21.04 0.58 0.0279
700 0.1365 76.8172 4.552021539 39.9306 259.1365 0.122715 0.078431 0.019402 -0.0813 12.30002 31.77339224 22.67 0.63 0.0300
720 0.1947 95.450817 3.310171498 4.953032 164.0128 0.114442 0.074294 0.027674 -0.21076 4.744841 32.26948281 23.03 0.64 0.0305
740 0.2530 117.22018 2.448923765 0.687754 106.4057 0.106158 0.070152 0.035958 -0.31354 3.189395 32.50072612 23.19 0.64 0.0307
760 0.3114 142.4063 1.840722016 0.105955 70.62234 0.097863 0.066004 0.044254 -0.38335 2.608548 32.66994025 23.31 0.65 0.0309
780 0.3698 171.28577 1.403975271 0.017967 47.86838 0.089558 0.061852 0.052559 -0.42636 2.345463 32.81469546 23.42 0.65 0.0310
800 0.4284 204.12856 1.08545581 0.003329 33.08263 0.081241 0.057694 0.060876 -0.43891 2.278364 32.94998863 23.51 0.65 0.0311
820 0.4869 241.19597 0.849798431 0.00067 23.27974 0.072914 0.05353 0.069202 -0.41257 2.423849 33.08773998 23.61 0.65 0.0313
840 0.5456 282.73889 0.673102847 0.000145 16.65804 0.064577 0.049361 0.07754 -0.3355 2.980587 33.24627371 23.72 0.66 0.0314
860 0.6044 328.9963 0.53895873 3.39E-05 12.10682 0.056228 0.045187 0.085888 -0.19205 5.207033 33.48675637 23.90 0.66 0.0316
872.02 0.6397 359.15994 0.473885696 1.46E-05 10.06454 0.051205 0.042675 0.090912 -0.06556 15.25393 33.84839374 24.15 0.67 0.0320
BED 1STOICHIOMETRIC TABLE
0
100
200
300
400
500
600
0.00 0.20 0.40 0.60 0.80
1/R
X(T)
Bed 1 Levenspiel Plot
Series1
Setting up the series of equations above on excel it is possible to find X(T) and the corresponding
rate at that instant. Then by plotting 1/R against X it is possible to determine the total weight of
catalyst required and the corresponding bed height and pressure drop, this is done by integrating
the resulting equation analytically using the trapezium rule. The results of each bed are shown
below:
Fig 1.5 – Showing the Stoichiometric Table and Derived Levenspiel Plot. In order to find the mass of catalyst required, the area under
the curve is multiplied by the feed molar flowrate of SO2into each Bed.
23. 22
0.00
5.00
10.00
15.00
20.00
25.00
30.00
0.00 0.10 0.20 0.30 0.40 0.50 0.60 0.70
Weight/MetricTonnes
Conversion X(T)
Bed 1 Catalyst Weight Vs Conversion
0.000
0.100
0.200
0.300
0.400
0.500
0.600
0.700
0.800
0.00 5.00 10.00 15.00 20.00 25.00 30.00
Height/m
Weight/ Metric Tonnes
Bed 1 Catalyst Weight Vs Height
Once the weight is known the height of catalyst packing in the bed is found by dividing the mass by
the bulk density of the catalyst packing and then dividing the resulting volume of catalyst bed is
divided by the cross-sectional area in order to determine the height of packing in the Bed and then
then corresponding pressure drop (refer to fig above). The following graphical was also derived from
this stoichiometric table.
32. 31
iv) Off-gas Cooling & Heating Requirements Between Beds (All
Temperatures in Kelvin)
The forward oxidation reaction occurring in the catalytic converter is exothermic, therefore there is
an operational and safety requirement to cool the gas leaving each catalyst bed. This is done by
modelling the heat exchange process as a system of variables in an array of vectors and solving
Once the excel model has data for enthalpy the multi-variable matrix calculations can be set up
based on the design basis and equilibrium constraints. The heat duty required to cool down or heat
up the off-gas can be determined by a similar matrix calculation to appendix 1i) with altered the
constraints on the goal seek solver. This is shown below:
Fig 1.1 – Bed 1 Outlet Off-gas cooling requirement. By specifying no change in composition during heat exchange, the heat loss
requirement to cool the outlet Off-gas can be calculated for any bed in the converter by numerical iteration using the goal seek tool.
By using the goal seek tool in the data window, I was able to set the difference in outlet and inlet
composition to zero by changing the numerical enthalpy term (heat loss) in the column vector, the
solution for heat removal duty requirement of a heat exchanger can be determined for any design
basis flow of off gas using this method. This method saves time and is numerically stable. It should
be noted that there is a degree of inaccuracy in the calculation due to truncation error as the
difference between the inlet and outlet temperature increases. This error is due to the linearized
enthalpy data used to calculate the enthalpies, this error can be removed by making use of the
thermal constants of specific heat capacity for each species. However since this error is so low
(roughly 1-5%), for ease of calculation the enthalpy data in table 1.1 was sufficient to use. It should
be noted that because my basis of flow was in Kg-mole per hour, the resulting heat removal duty is
in MJ per hour. To convert this heat duty to MJ per second I divided the value by 3600. For ease of
calculation the heat exchange process is assumed to be isobaric, and no energy loss through thermal
expansion of the off-gas occurs (although in a real or dynamic model this would not be the case). It
should also be noted that by setting the goal seek such that the inlet and outlet off gas molar flow
33. 32
Bed 3 Offgas Cooling
Numerical term SO3 in SO2 in O2 in CO2 in N2 in SO3 out SO2 out O2 out CO2 out N2 out
Feed SO3 kg-mole 678.8297688 1 0 0 0 0 0 0 0 0 0
Feed SO2 kg-mole 34.78298436 0 1 0 0 0 0 0 0 0 0
Feed O2 kg-mole 103.1202457 0 0 1 0 0 0 0 0 0 0
Feed CO2 kg-mole 90.21411332 0 0 0 1 0 0 0 0 0 0
Feed N2 kg-mole 4779.210253 0 0 0 0 1 0 0 0 0 0
S Balance 0 -1 -1 0 0 0 1 1 0 0 0
O Balance 0 -3 -2 -2 -2 0 3 2 2 2 0
C Balance 0 0 0 0 -1 0 0 0 0 1 0
N Balance 0 0 0 0 0 -2 0 0 0 0 2
Enthalpy Balance -50348.73261 368.9235 276.9676 -13.3122 374.5356 -12.6994 -386.309 -289.427 5.2036 -386.803 5.131
Inlet Gas Temp 723.3557
Outlet Temp 480
Heat Loss MJ/h -50348.7
SO3 in 678.8297688
SO2 in 34.78298436
O2 in 103.1202457
CO2 in 90.21411332
N2 in 4779.210253
SO3 out 678.8297688
SO2 out 34.78298436
O2 out 103.1202457
CO2 out 90.21411332
N2 out 4779.210253
Goal Finder 0.00000000
rates are the same, no reaction occurs therefore the enthalpy change is purely as a result of cooling
or heating up the gas. The results for each off gas heat exchange calculation is shown below:
Fig 1.2 – Bed 2 Outlet Off-gas cooling requirement. By specifying no change in composition during heat exchange, the heat loss
requirement to cool the outlet Off-gas can be calculated for any bed in the converter by numerical iteration using the goal seek tool.
Fig 1.3 – Bed 3 Outlet Off-gas cooling requirement. By specifying no change in composition during heat exchange, the heat loss
requirement to cool the outlet Off-gas can be calculated for any bed in the converter by numerical iteration using the goal seek tool.
Bed 2 Cooling
Numerical term SO3 in SO2 in O2 in CO2 in N2 in SO3 out SO2 out O2 out CO2 out N2 out
Feed SO3 kg-mole 629.8262346 1 0 0 0 0 0 0 0 0 0
Feed SO2 kg-mole 83.78651854 0 1 0 0 0 0 0 0 0 0
Feed O2 kg-mole 127.6220128 0 0 1 0 0 0 0 0 0 0
Feed CO2 kg-mole 90.21411332 0 0 0 1 0 0 0 0 0 0
Feed N2 kg-mole 4779.210253 0 0 0 0 1 0 0 0 0 0
S Balance 0 -1 -1 0 0 0 1 1 0 0 0
O Balance 0 -3 -2 -2 -2 0 3 2 2 2 0
C Balance 0 0 0 0 -1 0 0 0 0 1 0
N Balance 0 0 0 0 0 -2 0 0 0 0 2
Enthalpy Balance -14542.92053 365.564 274.5407 -14.8791 372.1651 -14.1618 -370.592 -278.073 12.534 -375.713 11.973
Inlet Gas Temp 770.3803
Outlet Temp 700
Heat Loss MJ -14542.9
SO3 in 629.8262346
SO2 in 83.78651854
O2 in 127.6220128
CO2 in 90.21411332
N2 in 4779.210253
SO3 out 629.8262346
SO2 out 83.78651854
O2 out 127.6220128
CO2 out 90.21411332
N2 out 4779.210253
Goal Finder 0.000000000
34. 33
Fig 1.4 – Bed 4 Inlet Off-gas heating requirement. By specifying no change in composition during heat exchange, the heat loss
requirement to cool the outlet Off-gas can be calculated for any bed in the converter by numerical iteration using the goal seek tool.
The positive sign (green highlight) indicates that heat is being supplied to heat the gas from 300k to 683k (which is logical).
The data above has been used to design 4 shell and tube heat exchangers following guidelines
specified in Coulson and Richardson Volume 6 for the operating conditions specified in my basis of
design and plant specifications (see plant specification sheet for details)4
.
Fig 1.5 – Relationship between overall Heat Transfer Co-efficient U and the shell side and tube side heat transfer co-efficients of various
process and service fluids. I used this graph as a starting point for the design of my heat exchangers and then iterated the design
parameters to find the actual value of U in each heat exchanger using the goal finder tool. This process is then used to optimise the
heat exchanger to minimise the capital costs of each heat exchanger.
4
C&R Vol 6 4th
Edition Chapter 12
Bed 4 Inlet Gas Heating After Interpass Absorber
Numerical term SO3 in SO2 in O2 in CO2 in N2 in SO3 out SO2 out O2 out CO2 out N2 out
Feed SO3 kg-mole 13.57659538 1 0 0 0 0 0 0 0 0 0
Feed SO2 kg-mole 34.78298436 0 1 0 0 0 0 0 0 0 0
Feed O2 kg-mole 103.1202457 0 0 1 0 0 0 0 0 0 0
Feed CO2 kg-mole 90.21411332 0 0 0 1 0 0 0 0 0 0
Feed N2 kg-mole 4779.210253 0 0 0 0 1 0 0 0 0 0
S Balance 0 -1 -1 0 0 0 1 1 0 0 0
O Balance 0 -3 -2 -2 -2 0 3 2 2 2 0
C Balance 0 0 0 0 -1 0 0 0 0 1 0
N Balance 0 0 0 0 0 -2 0 0 0 0 2
Enthalpy Balance MJ/h 53077.73782 395.596 296.2365 -0.872 393.3565 -1.088 -371.806 -278.95 11.96756 -376.57 11.4443
Feed Gas Temp 350
Product Temp 683
Heat Required MJ/h 53077.74
Catalyst Bed 2 Composition
SO3 in 13.57659538
SO2 in 34.78298436
O2 in 103.1202457
N2 in 90.21411332
CO2 in 4779.210253
SO3 out 13.57659538
SO2 out 34.78298436
O2 out 103.1202457
CO2 out 90.21411332
N2 out 4779.210253
Goal Finder 0.000000000
35. 34
Based On The Calculations carried out for each parameter of the shell and tube design (fig 1.6-1.9) I
obtained values for the shell and tube design. I then carried out a mechanical design to obtain design
parameters such as the shell thickness and saddle support thicknesses. These have been omitted in
this report as these are not required to generate a cost estimate of the heat exchanger. The
manufacturer of a heat exchanger will often specify this information in the material specification
sheet that comes with the purchase.
I used steam tables to determine the dryness fraction of the steam generated in each stage. The
enthalpy of the mixture at any point is equal to:
ℎ 𝑚1 = ℎ𝑙1 + 𝑥ℎ 𝑔1
∴ ∆ℎ 𝑚 = ℎ 𝑚2 − ℎ 𝑚1
Where ∆ℎ 𝑚 is the specific enthalpy change in the offgas mixture or the change in specific enthalpy
of the service fluid (found using steam tables) at a given temperature. This is amount of heat that
needs to be absorbed by the service fluid to reduce to temperature of the offgass. The heat given off
by the offgas during the gas cooling is equal to the heat absorbed by the service fluid. Therefore by
dividing the heat duty requirement Q by ∆ℎ 𝑚 , the mass flowrate of the cooling water required is
found. Using the goal finder tool the dryness fraction of steam in each bed can be found and
therefore the specific enthalpy of the service fluid at any operating temperature and pressure in the
design. It should be noted that the entropy change in the fluid causes the vapour to condense as the
pressure is increased from heat exchanger 1 to 2. The additional heat supplied in heat exchangers 2
& 3 is sufficient to produce 34.6 MT/h of saturated steam at 35bar g at a 0.95 dryness fraction.
36. 35
Q GJ/h -37.9 Pressure Drop Calculation
hv (GJ/MT) 2.7403 Np 1
hm (GJ/MT) 0.105215411 m -0.14
hl (GJ/MT) 0.602100028 jf 2.00E-03
Dryness Fraction x 0.218440336 l/di 108.695652
Δhm 1.095476671 μ/μw 1.2
m (MT/h) 34.61556578 c 2.5
m (kg/s) 9.615434939 ρ 0.58979054
molar flow kgmol/s 1.610267746 ut 160.696828
mass flow kg/s 54.92499487 ΔP (bar) 1.99E-03
volumetric flow rate m^3/s 92.17233114 Tube Side Outlet Pressure 1.2
A (m^2) 249.3821887 1.013
outer diameter(m) 0.05 Tube Pump Pressure Rating 2.01E-01
thickness (m) 0.002 jf 2.50E-03
tube-side velocity (m/s) 160.6968276 m -0.14
Total Pipe Length (m) 1725.44619 Ds/de 6.74816956
Individual Pipe Length (m) 5 L/Lb 25
Number of Tube Passes 1 ρ 958
Number of Pipes 345.089238 us 0.02389759
Gas Viscosity 5.50E-05 ΔP (bar) 3.98E-02
Gas Density 0.58979054 Shell Side Outlet Pressure 4.013
cp gas mixture 1068 Suction Side Pressure (bar) 4.103
Thermal Conductivity 0.0615 Shell Pump Pressure Rating (bar) 3.51E-01
Reynold's Number Re 8.62E+04
Prandtl Number Pr 9.55E-01
Heat Transfer Factor 2.00E-03
Nusselt Number Nu 1.70E+02
hi 2.09E+02
molar flow kgmol/s 0.53419083
mass flow kg/s 9.615434939
volumetric flow rate m^3/s 0.010036988
Shell Diameter m 3
Tube Pitch m 0.15
Baffle Cut 0.2
Baffle Spacing 0.21
Total Shell Side Area m^2 10.5
A (m^2) Per Baffle Pass 0.42
outer diameter(m) 0.05
thickness (m) 0.002
Effective Diameter 0.444565
shell-side velocity (m/s) 0.023897592
Shell Length (m) 5.25
Number of Baffle Spaces 25
Number of Shell Passes 1
Cooling Fluid Viscosity 0.0003
Fluid density 958
cp fluid 3708.397223
Thermal Conductivity 0.7
Reynold's Number Re 3.39E+04
Prandtl Number Pr 1.589313096
Heat Transfer Factor 2.50E-03
Nusselt Number Nu 9.88E+01
ho 3.29E+02
kw (SS-Stainless Steel) 45
hod 500
do 0.05
di 0.046
hid 5000
Trial U (W/m^2/K) 120.6372804
dt (K) -350.1265922
Calculated U (W/m^2/K) 1.21E+02
Goal Finder 0.0000
Tube Side
Pressure Drop
Shell Side
Pressure Drop
Suction Side Pressure (bar)
Shell Side
Heat Transfer
Co-efficient
Calculation
and Tube Side
Surface Area
Requirement
Material & Dirt
Co-efficients
Tube Side
Heat Transfer
Co-efficient
Calculation
and Tube Side
Surface Area
Requirement
Mass Flow
Rate of
Cooling Liquid
(3 bar g, 4.013
bar abs)
Bed 1 Heat Exchanger (Boiler)
HEAT EXCHANGER 1 SPECIFICATION
37. 36
Q GJ/h -14.5 Pressure Drop Calculation
hv (GJ/MT) 2.6034 Np 1
hm 1.200692082 m -0.14
hl 1.05669 jf 2.00E-03
Dryness Fraction x 0.21668921 l/di 86.95652174
Δhm 0.420126616 μ/μw 1.2
m (MT/h) 34.61556578 c 2.5
m (kg/s) 9.615434939 ρ 0.674514695
molar flow kgmol/s 1.610267746 ut 208.005597
mass flow kg/s 54.92499487 ΔP (bar) 2.71E-03
volumetric flow rate m^3/s 81.42890774 Tube Side Outlet Pressure 1.2
A (m^2) 136.1650781 Suction Side Pressure (bar) 1.2
outer diameter(m) 0.05 Tube Pump Pressure Rating 1.47E-02
thickness (m) 0.002 jf 2.50E-03
tube-side velocity (m/s) 208.005597 m -0.14
Total Pipe Length (m) 942.1102459 Ds/de 6.748169559
Individual Pipe Length (m) 4 L/Lb 20
Number of Tube Passes 1 ρ 750.6717605
Number of Pipes 235.5275615 us 0.030497874
Gas Viscosity 5.50E-05 ΔP (bar) 3.19E-02
Gas Density 0.674514695 Shell Side Outlet Pressure 36.013
cp gas mixture 1068 Suction Side Pressure (bar) 4.013
Thermal Conductivity 0.0615 Shell Pump Pressure Rating 3.56E+01
Reynold's Number Re 1.28E+05
Prandtl Number Pr 9.55E-01
Heat Transfer Factor 2.00E-03
Nusselt Number Nu 2.51E+02
hi 3.09E+02
molar flow kgmol/s 0.53419083
mass flow kg/s 9.615434939
volumetric flow rate m^3/s 0.012809107
Shell Diameter m 3
Tube Pitch m 0.15
Baffle Cut 0.2
Baffle Spacing 0.21
Total Shell Side Area m^2 8.4
A (m^2) Per Baffle Pass 0.42
outer diameter(m) 0.05
thickness (m) 0.002
Effective Diameter 0.444565
shell-side velocity (m/s) 0.030497874
Shell Length (m) 4.2
Number of Baffle Spaces 20
Number of Shell Passes 1
Cooling Fluid Viscosity 0.00012
Fluid density 750.6717605
cp fluid 3712.233933
Thermal Conductivity 0.7
Reynold's Number Re 8.48E+04
Prandtl Number Pr 0.63638296
Heat Transfer Factor 2.50E-03
Nusselt Number Nu 1.83E+02
ho 6.09E+02
kw (SS-Stainless Steel) 45
hod 500
do 0.05
di 0.046
hid 5000
Trial U (W/m^2/K) 187.2011986
dt (K) -158.4801106
Calculated U (W/m^2/K) 1.87E+02
Goal Finder 0.0000
Tube Side
Pressure
Drop
Shell Side
Pressure
Drop
Tube Side Heat
Transfer Co-
efficient
Calculation and
Tube Side Surface
Area
Requirement
Shell Side Heat
Transfer Co-
efficient
Calculation and
Tube Side Surface
Area
Requirement
Material & Dirt Co-
efficients
Mass Flow Rate
of Cooling Liquid
(4.013 bar
Saturated Liquid
to 36.013 bar
Saturated Liquid)
x=0.0
Bed 2 Heat Exchanger (Boiler)
HEAT EXCHANGER 2 SPECIFICATION
38. 37
Q GJ/h -50.34873261
hv (GJ/MT) 2.6034 Np 2
hm (GJ/MT) 1.620818699 m -0.14
hl (GJ/MT) 0.602100028 jf 0.002
Dryness Fraction x 0.95 l/di 130.4347826
Δhm (GJ/MT) 1.45451133 μ/μw 1.2
m (MT/h) 34.61556578 c 2.5
m (kg/s) 9.615434939 ρ 0.674514695
molar flow kgmol/s 1.610267746 ut 377.3125354
mass flow kg/s 54.92499487 ΔP (bar) 0.011540078
volumetric flow rate m^3/s 81.42890774 Tube Side Outlet Pressure (Bar) 1.2
A (m^2) 225.1960565 Suction Side Pressure 1.2
outer diameter(m) 0.05 Tube Pump Pressure Rating 0.023540078
thickness (m) 0.002 jf 0.0025
tube-side velocity (m/s) 377.3125354 m -0.14
Total Pipe Length (m) 1558.10517 Ds/de 4.498779706
Individual Pipe Length (m) 6 L/Lb 30
Number of Tube Passes 2 ρ 1.2
Number of Pipes 129.8420975 us 0.700261808
Gas Viscosity 5.50E-05 ΔP (bar) 0.001170106
Gas Density 0.674514695 Shell Side Outlet Pressure (Bar) 36.013
cp gas mixture 1068 Suction Side Pressure 36.013
Thermal Conductivity 0.0615 Shell Pump Pressure Rating (Bar) 3.602470106
Reynold's Number Re 2.31E+05
Prandtl Number Pr 9.55E-01
Heat Transfer Factor 0.002
Nusselt Number Nu 455.7738174
hi 560.6017954
molar flow kgmol/s 0.53419083
mass flow kg/s 9.615434939
volumetric flow rate m^3/s 0.196073306
Shell Diameter m 2
Tube Pitch m 0.15
Baffle Cut 0.2
Baffle Spacing 0.21
Total Shell Side Area m^2 16.8
A (m^2) Per Baffle Pass 0.28
outer diameter(m) 0.05
thickness (m) 0.002
Effective Diameter 0.444565
shell-side velocity (m/s) 0.700261808
Shell Length (m) 6.3
Number of Baffle Spaces 30
Number of Shell Passes 2
Cooling Fluid Viscosity 0.0003
Fluid density 49.04
cp fluid 2105.55
Thermal Conductivity 0.7
Reynold's Number Re 68365.00263
Prandtl Number Pr 0.902378571
Heat Transfer Factor 0.0025
Nusselt Number Nu 165.2160135
ho 550.720045
kw (SS-Stainless Steel) 45
hod 500
do 0.05
di 0.046
hid 5000
Trial U (W/m^2/K) 444.9365509
dt (K) -139.5812802
Calculated U (W/m^2/K) 444.9364897
Goal Finder 0.000
Bed 3 Heat Exchanger (Boiler)
Pressure Drop Calculation
Shell
Side
Pressure
Drop
Mass Flow Rate
of Cooling
Liquid (35 bar g,
36.013 bar abs)
Tube Side Heat
Transfer Co-
efficient
Calculation and
Tube Side
Surface Area
Requirement
Tube
Side
Pressure
Drop
Shell Side Heat
Transfer Co-
efficient
Calculation and
Tube Side
Surface Area
Requirement
Material & Dirt
Co-efficients
HEAT EXCHANGER 3 SPECIFICATION
39. 38
Q GJ/h 68.5808108
hv (GJ/MT) 2.6034 Np 2
hm (GJ/MT) 3.07533003 m -0.14
hl (GJ/MT) 0.3796 jf 0.002
Dryness Fraction x 0 l/di 108.6956522
Δhm (GJ/MT) -3.45493003 μ/μw 1.2
m (MT/h) 19.8501302 c 2.5
m (kg/s) 5.51392505 ρ 0.71836414
molar flow kgmol/s 1.61026775 ut 243.3229571
mass flow kg/s 54.9249949 ΔP (bar) 0.007333155
volumetric flow rate m^3/s 76.4584308 Tube Side Outlet Pressure (Bar) 1.2
A (m^2) 273.240112 Suction Side Pressure 1.2
outer diameter(m) 0.05 Tube Pump Pressure Rating 0.019333155
thickness (m) 0.002 jf 0.0025
tube-side velocity (m/s) 243.322957 m -0.14
Total Pipe Length (m) 1890.51637 Ds/de 56.33406098
Individual Pipe Length (m) 5 L/Lb 30
Number of Tube Passes 2 ρ 479.6
Number of Pipes 189.051637 us 0.16424178
Gas Viscosity 0.000055 ΔP (bar) 1.373480616
Gas Density 0.71836414 Shell Side Outlet Pressure 1.013
cp gas mixture 1068 Suction Side Pressure 36.103
Thermal Conductivity 0.0615 Shell Pump Pressure Rating 1.373480616
Reynold's Number Re 158904.079
Prandtl Number Pr 0.95512195
Heat Transfer Factor 0.002
Nusselt Number Nu 313.028913
hi 385.025564
molar flow kgmol/s 0.30632917
mass flow kg/s 5.51392505
volumetric flow rate m^3/s 0.01149692
Shell Diameter m 2
Tube Pitch m 0.0625
Baffle Cut 0.2
Baffle Spacing 0.175
Total Shell Side Area m^2 4.2
A (m^2) Per Baffle Pass 0.07
outer diameter(m) 0.05
thickness (m) 0.002
Effective Diameter 0.0355025
shell-side velocity (m/s) 0.16424178
Shell Length (m) 5.25
Number of Baffle Spaces 30
Number of Shell Passes 2
Cooling Fluid Viscosity 0.00012
Fluid density 479.6
cp fluid 4187
Thermal Conductivity 0.7
Reynold's Number Re 196017.914
Prandtl Number Pr 0.71777143
Heat Transfer Factor 0.0025
Nusselt Number Nu 439.249431
ho 1756.99772
kw (SS-Stainless Steel) 45
hod 500
do 0.05
di 0.046
hid 5000
Trial U (W/m^2/K) 275.474897
dt (K) -253.08921
Calculated U (W/m^2/K) 275.474712
Goal Finder 0.000
Bed 4 Feed Gas Heater (Steam Condenser)
Pressure Drop Calculation
Shell
Side
Pressure
Drop
Mass Flow Rate of
Cooling Liquid
(36.013 bar
Saturated Steam to
1.013 bar Water at
355K) x=0.0
Tube Side Heat
Transfer Co-
efficient
Calculation and
Tube Side Surface
Area Requirement
Shell Side Heat
Transfer Co-
efficient
Calculation and
Tube Side Surface
Area Requirement
Material & Dirt Co-
efficients
Tube
Side
Pressure
Drop
HEAT EXCHANGER 4 SPECIFICATION
41. 40
19.25266772 0.996159055
11 55.44800083
£1,484,149.65 88
£29,920,628.42
Qb Q Cb /£
Cost
Component
M
Non Inflatted FOB
Price Cr Base Year
Plant Cost
Inflation from
2004-2013
Inflatted FOB Price
CCe Pressure Factor fp
Material Factor
fm
Temperature
Factor ft
Equipment
Cost Ce
Blower (Base Unit =kW) 20 3.91E+02 1160 0.8 £12,515.01 2004 1.23813 £15,495.25 1.0 1.0 1.0 £16,270.01
Auxillary Blower (Base Unit =kW) 20 3.88E+02 1160 0.8 £12,432.68 2004 1.23813 £15,393.32 1.0 1.0 1.0 £16,162.99
Primary Pump 1(Base Unit =kW) 20 1.70E+02 1160 0.8 £6,438.40 2004 1.23813 £7,971.59 1.0 1.0 1.0 £8,370.17
Auxillary Pump 1(Base Unit =kW) 20 1.70E+02 1160 0.8 £6,438.40 2004 1.23813 £7,971.59 1.0 1.0 1.0 £8,370.17
Primary Pump 2(Base Unit =kW) 20 1.19E+02 1160 0.8 £4,840.12 2004 1.23813 £5,992.71 1.0 1.0 1.0 £6,292.35
Auxillary Pump 2(Base Unit =kW) 20 1.19E+02 1160 0.8 £4,840.12 2004 1.23813 £5,992.71 1.0 1.0 1.0 £6,292.35
Primary Pump 3(Base Unit =kW) 20 8.35E+01 1160 0.8 £3,638.60 2004 1.23813 £4,505.08 1.0 1.0 1.0 £4,730.33
Auxillary Pump 3(Base Unit =kW) 20 8.35E+01 1160 0.8 £3,638.60 2004 1.23813 £4,505.08 1.0 1.0 1.0 £4,730.33
HE1(Base Unit =ft) 817.9736 £8,250.00 2003 1.34131 £11,065.82 1.0 2.0 1.8 £41,828.80
HE2(Base Unit =ft) 446.6215 £5,940.00 2003 1.34131 £7,967.39 1.5 2.0 1.4 £35,136.19
HE3(Base Unit =ft) 738.6431 £7,590.00 2003 1.34131 £10,180.55 1.5 2.0 1.4 £44,896.24
HE4(Base Unit =ft) 693.6303 £8,580.00 2003 1.34131 £11,508.45 1.0 2.0 1.4 £33,834.85
Converter Packing
Bed 1(Base Unit =kg) £287,440.28 2013 1.00000 £287,440.28 1.0 1.0 1.0 £301,812.29
Bed 2(Base Unit =kg) £536,422.27 2013 1.00000 £536,422.27 1.0 1.0 1.0 £563,243.39
Bed 3(Base Unit =kg) £621,159.03 2013 1.00000 £621,159.03 1.0 1.0 1.0 £652,216.98
Bed 4(Base Unit =kg) £385,930.10 2013 1.00000 £385,930.10 1.0 1.0 1.0 £405,226.60
Converter Shell +Internal Fittings 421191.72 £297,000.00 2003 1.34131 £398,369.50 1.0 2.4 1.8 £1,807,004.07
PCE £3,956,418.12
£31,404,778.07
Estimated Inventory Value /Yr
Price (£/MT)
Annual Revenue From Sale (£/Yr)
Total SO3Production Rate (MT/h)
Market Price (£/MT)
Inventory Value £/Yr (Estimate)
Excess Steam From Superheater (MT/h) Overall Conversion Achieved After Bed 4
Market Value
Major
Equipment
Cost (PCE) £
Battery Limit Cost
Packing Cost Estimate =£11.90per kg Packing
Caesium Enriched Courtesy BASF
Data From Correlation
Chart Sneider Et Al 2003
Data From Correlation
Chart Sneider Et Al 2003
MATERIAL COSTING SHEET & OUTPUT INVENTORY
Fig 1.2 – Estimated Reactor battery limit fixed costs. PCE = £3,956,418.12
£/Yr
42. 41
BUILDING MATERIAL & ERECTION COSTS
Plant Battery = Limit Erection Cost
The capital cost of 9.5 million pounds covers the cost of manufacturing the equipment and
building the plant itself. The physical plant cost of the overall process should not come to
more than double this amount, provided the converter is the most expensive piece of
equipment in the plant, which is the case. The operating costs have not been considered in
this report as it is better practice to consider the total operating capital requirement of the
system and not the battery region of the converter.
For the investment to be considered tangible by investors, the rate of return on the Net
Present Value of the Capital Investment has to be atleast 20% per annum to fulfil a payback
period of 5 years. With an estimated revenue generated of 31 million pounds per annum
based on the inventory value of the saleable chemical products produced at the end of the
catalytic reaction, in order to meet this quota the plant would have to make 4 million pounds
a year in operating profit, which is a very achievable sales target based on these figures.
Fig 3.1 – Trade-off between fixed capital requirement of battery region and potential annual
earnings based on the NPV of inventory chemicals and saleable commodities produced during
catalytic conversion in the double contact process.
Factored Cost Estimates To Errect Plant Battery Region
Fixed Cost Factor
Equipment Erection 0.4
Piping 0.7
Instrumentation 0.3
Electrical 0.1
Total 1.5
PPC £5,934,627.18
Contractors Fee 0.05
Design + Engineering 0.15
Contingency 0.1
Total 0.3
Fixed Capital £9,495,403.49
£9,500,000.00
(Battery Region Fixed Capital)
£31,000,000.00
(Estimated Minimum Annual
Revenue)
43. 42
Appendix 4) Safety Considerations &
Hazop
START-UP PROCEDURE
i) Cool air is fed into the fired gas heat exchanger, and warmed up to 800-850
Kelvins.
ii) Valves 02-CV-01,02,03,04,05 and 06 are shut at this point. This allows warm
air to flow radially in beds 1-4 through the central pipeline and pressurize
the vessel bed to 1.2bar.
iii) Valves 02-CV-01,02,03,04,05 and 06 remain shut until the differential
pressure between the top and the bottom beds is equal to 0.1Po where
Po=Operating Pressure of 1.2bar. At this point, these valves will open
accordingly. Feedback Forward & Feedback control are used to determine
the fraction that each valve shuts or closes by. In order for the control
response to be fast (i.e. the nth order time constant to be low), electrical
signal processing is used to actuate the valves based on dynamic
pressure, temperature and inlet flow rate values.
iv) This signal sent instantly to the control room and will show up on the
system monitors. Manual override commands can be performed on each
control valve at any time as an additional safety feature.
v) At Po=1.2bar dynamic gas flow in and out of each bed is controlled by
microprocessors within the electronic control circuit which work based on
specified constraints set using a Laplace domain PID controller.
vi) The time taken for each bed to reach steady state is aimed to be 45-70
minutes depending on climate conditions. However a 3 dimensional CFD
model or Fourier heat model (or a scaled down model of reactor prototype)
would be required to determine the exact time requirement, an accurate
software of choice would be COMSOL Multiphysics or MATLAB (Simulink).
vii) Once the bed is at steady state, the flow of air through the fired gas heat
exchanger is stopped and the gate valve shut controlling air flow in is shut
and cool off gas is fed into the waste heat boiler from the main process
and then to the sulphur burner to enrich the stream with Sulphur Dioxide.
viii) Start-up is now complete, valves 04-CV-01, 02, 03 and 04 are shut fully.
SHUTDOWN PROCEDURE
i) Essentially the same process, no need for pressurization and pressure
maintenance of the vessel at 1.2bar. The fired gas heat exchanger is
bypassed and dry air is simply passed through the beds and out through
the bed outlets allowing cooling to take place.
AS A SAFETY FEATURE THE VESSEL CAN BE VENTED IN THE EVENT OF AN
EMERGENCY IF RAPID REMOVAL OF OFF GAS IS REQUIRED. THERE IS ALSO A
PRESSURE RELIEF VALVE SET TO BURST AT 2.4 BAR ON THE TOP OF THE
CONVERTER. A HIGH LEVEL ALARM WILL SOUND IF EITHER OF THESE SAFETY
FEATURES ARE USED.
44. 43
Appendix I) Safety Considerations & Hazop
HAZOP STUDY OF CONVERTER UNIT BATTERY LIMIT CONSIDERING EXTERNAL UPSTREAM AND DOWNSTREAM FACTORS
Line Number Guide Word Element Deviation Possible Cause Consequence Safeguards Actions & Recommendations Actions Assigned to
200-100HS01-FG-P1 No Process
Line Piping
&
Equipment
Flow Pipe Blockage Upstream Blower Damage Actuated Primary (PRIM) & Auxiliary
Back Up Pump (AUX) Lines
Low Level Alarm On Primary Line
Failure
Line Manager -
Control Room
Pump Failure Pipeline Burstage Upstream
01-FV-01 Failure Plant-Shutdown Flow Indicators On PRIM & AUX Lines High Level Alarm & Emergency
Shutdown On Failure Of Back
Pump (B-01-AUX)
B-01-PRIM Failure Loss Of Production Time Manual Gate Valve On AUX & AUX
LinesIncurred Operating Loss
As Well As Process
Line Piping
&
Equipment
Contaminants
& Debris
Contaminants Debris or Dirt
In Upstream Offgas
Blower Impeller Damage
Filter At Off Gas Inlet Attached To High
Level Flow Indicator & Alarm
Regular Pipe Maintenance &
Replacement Of Corroded
Piping, Valves & Equipment
Line Manager -
Action By Line
Technician
Build-up Of Debris In Piping
Rust From Corroded Piping &
Equipment Upstream/On
Process Line Product Contamination
Down Stream Equipment &
Piping Failure
Process
Line Piping
&
Equipment
Water Vapour Insufficient Gas Drying
Mist Formation In Beds Demister Pads At Bed Inlet & Outlet Ensure Inlet Offgas is dry
Process Engineer &
Control EngineerHigh Initial water
concentration in off-gas Corrosion In Pipe Gas Drying Before Line Ensure Air Used To Heat Up Beds
During Start-up is dryConverter Shell Corrosion
Less Process
Operating
Condition
Temperature Excess Gas Cooling after
Sulphur Burner
Reduced Bed Conversion Bypass Line At heat exchanger 1 For
temperature control
Bypass fraction of process fluid
around heat exchanger 1
Line Manager -
Action By Line
Technician
Reduced Bed Temperature
More Process
Operating
Condition
Temperature Insufficient Gas Cooling after
Suphur Burner
Increased Bed Temperature Temperature Control via Service
Stream Flow Control
Increase Cooling Water Flow
RateReduced Bed Efficiency
Structural Damage
200-100HS01-FG-P2 Less Process
Operating
Condition
Temperature
Fouling In Catalyst Bed 1 Reduced Bed Conversion
Bypass Line At Heat Exchanger 2 For
Temperature Maintenance
Bypass fraction of process fluid
around heat exchanger 2 Line Manager -
Action By Line
Technician
Excess Gas Cooling after
Sulphur Burner Reduced Bed Temperature
Denatured Catalyst Packing in
Bed 1
Temperature Indicator & Controller on
Line
Ensure Catalyst bed is not
heated above 900K
More Process
Operating
Condition
Temperature Insufficient Gas Cooling after
Sulphur Burner Increased Bed Temperature
Temperature Control via Service
Stream Flow Control Increase Cooling Water Flow
Rate
Line Manager -
Action By Line
Technician
Reduced Bed Efficiency
Structural Damage
No Process
Line Piping
&
Equipment
Flow Pipe blockage in 200-
100HS01-FG-P1
Loss Of Production Time Central Bypass Line (Bed 1 and line
200-100HS01-FG-P2 can be bypassed
through central pipeline)
Bypass Line and Replace Blocked
Piping Line Manager -
Control Room
Incurred Operating Loss
Pipe Blockage In Stream Plant-Shutdown
45. 44
As Well As Process
Line Piping
&
Equipment
Contaminants
& Debris
Contaminants Debris or Dirt
In Upstream Offgas
Build-up Of Debris In Piping
Filter At Off Gas Inlet Attatched To
High Level Flow Indicator & Alarm
Regular Pipe Maintenance &
Replacement Of Corroded
Piping, Valves & Equipment
Line Manager -
Action By Line
Technician
Product Contamination
Rust From Corroded Piping &
Equipment Upstream/On
Process Line Down Stream Equipment &
Piping Failure
Process
Line Piping
&
Equipment
Water Vapour Insufficient Gas Drying Mist Formation In Beds Demister Pads At Bed Inlet & Outlet Ensure Inlet Offgas is dry
Process Engineer &
Control Engineer
High Initial water
concentration in off-gas Corrosion In Pipe Gas Drying Before Line Ensure Air Used To Heat Up Beds
During Start-up is dryConverter Shell Corrosion
More Process
Line Piping
&
Equipment
Flow Increased Off Gas Feed Flow
Rate
Increased Bed Temperature Temperature Indicator & Controller on
Line
Ensure Feedback Loop Is
Properly Calibrated To Control
OFF-GAS temperature in line
Line Manager -
Action By Line
Technician
Reduced Bed Conversion
Increased Sulphur Burning In
Burner
Potential Pipe Burstage Feed Back Loop to Line 40-100HC01-
SW-S1 flow Controller
200-100HS01-FG-P3 Less Process
Operating
Condition
Temperature Fouling In Catalyst Bed Reduced Bed 2 Conversion Bypass Line At Heat Exchanger 2 For
Temperature Maintenance
Bypass process fluid around heat
exchanger 1 Line Manager -
Action By Line
Technician
Excess Gas Cooling after
Sulphur Burner Reduced Bed Temperature
Denatured Catalyst Packing
Temperature Indicator & Controller on
Line
Ensure Catalyst bed is not
heated above 900K
More Process
Operating
Condition
Temperature Increased temperature in FG-
P2 Increased Bed 2 Temperature
Feed Back Loop to Line 40-100HC01-
SW-S1 flow Controller
Increase Cooling Water Flow
Rate
Line Manager -
Action By Line
Technician
Reduced Bed Efficiency
Structural Damage
No Process
Line Piping
&
Equipment
Flow Pipe blockage in 200-
100HS01-FG-P2
Loss Of Production Time Central Bypass Line (Bed 2 and line
200-100HS01-FG-P3 can be bypassed
through central pipeline)
Bypass Line and Replace Blocked
Piping Line Manager -
Control RoomIncurred Operating Loss
Pipe Blockage In Stream Plant-Shutdown
As Well As Process
Line Piping
&
Equipment
Contaminants
& Debris
Contaminants Debris or Dirt
In Upstream Offgas
Build-up Of Debris In Piping
Filter At Off Gas Line Inlet Attached To
High Level Flow Indicator & Alarm
Regular Pipe Maintenance &
Replacement Of Corroded
Piping, Valves & Equipment
Line Manager -
Action By Line
Technician
Offgas Contamination
Rust From Corroded Piping &
Equipment Upstream/On
Process Line
Down Stream Equipment &
Piping Failure
Process
Line Piping
&
Equipment
Water Vapour Insufficient Gas Drying Mist Formation In Beds Demister Pads At Bed Inlet & Outlet Ensure Inlet Offgas is dry
Process Engineer &
Control Engineer
High Initial water
concentration in off-gas Corrosion In Pipe Gas Drying Before Line Ensure Air Used To Heat Up Beds
During Start-up is dryConverter Shell Corrosion
More Process
Line Piping
&
Equipment
Flow Increased Off Gas Feed Flow
Rate
Increased Bed 2 Temperature Temperature Indicator & Controller on
Line
Ensure Feedback Loop Is
Properly Calibrated To Control
OFF-GAS temperature in line
Line Manager -
Action By Line
Technician
Reduced Bed 2 Conversion
Increased Sulphur Burning In
Burner
Potential Pipe Burstage Feed Back Loop to Line 40-100HC01-
SW-S1 flow Controller
46. 45
200-100HS01-FG-P4 Less Process
Operating
Condition
Temperature Fouling In Catalyst Bed 2 Increased Bed 3 Temperature Bypass line at heat exchanger 3 for
Temperature Maintenance
Bypass fraction of process fluid
around heat exchanger 2 Line Manager -
Action By Line
Technician
Excess Gas Cooling after
Sulphur Burner Reduced Bed 3 conversion
Denatured Catalyst Packing
Temperature Indicator & Controller on
Line
Ensure Catalyst bed is not
heated above 900K
More Process
Operating
Condition
Temperature Temperature increase in Bed
2 Increased Bed 3 Temperature
Feed Back Loop to Line 40-100HC01-
SW-S1 flow Controller
Increase Cooling Water Flow
Rate
Line Manager -
Action By Line
Technician
Reduced Bed 3 Conversion
Structural Damage
No Process
Line Piping
&
Equipment
Flow pipe blockage in 200-
100HS01-FG-P3
Loss Of Production Time Central Bypass Line (Bed 3 and line
200-100HS01-FG-P4&5 can be
bypassed through central pipeline)
Bypass Line and Replace Blocked
Piping
Line Manager -
Control Room
Incurred Operating Loss
Pipe Blockage In Stream Plant-Shutdown
As Well As Process
Line Piping
&
Equipment
Contaminants
& Debris
Contaminants Debris or Dirt
In Upstream Offgas
Buildup Of Debris In Piping Filter At Off Gas Line Inlet Attatched
To High Level Flow Indicator & Alarm
Regular Pipe Maintenance &
Replacement Of Corroded
Piping, Valves & Equipment
Line Manager -
Action By Line
Technician
Product Contamination
Rust From Corroded Piping &
Equipment Upstream/On
Process Line
Down Stream Equipment &
Piping Failure
Process
Line Piping
&
Equipment
Water Vapour Insufficient Gas Drying Mist Formation In Beds Demister Pads At Bed Inlet & Outlet Ensure Inlet Offgas is dry
Process Engineer &
Control Engineer
High Initial water
concentration in off-gas Corrosion In Pipe Gas Drying Before Line Ensure Air Used To Heat Up Beds
During Start-up is dryConverter Shell Corrosion
More Process
Line Piping
&
Equipment
Flow Increased Off Gas Feed Flow
Rate
Increased Bed Temperature Temperature Indicator & Controller on
Line
Ensure Feedback Loop Is
Properly Calibrated To Control
OFF-GAS temperature in line
Line Manager -
Action By Line
Technician
Reduced Bed Conversion
Increased Sulphur Burning In
Burner
Potential Pipe Burstage Feed Back Loop to Line 40-100HC01-
SW-S1 flow Controller
200-100HS01-FG-P5 Less Process
Operating
Condition
Temperature Fouling In Catalyst Bed Increased Bed Temperature Temperature Indicator & Controller on
Line
Bypass process fluid around heat
exchanger 1 Line Manager -
Action By Line
Technician
Excess Gas Cooling after
Sulphur Burner Reduced Bed Conversion
Denatured Catalyst Packing
Temperature Indicator & Controller on
Line
Ensure Catalyst bed is not
heated above 900K
More Process
Operating
Condition
Temperature Insufficient Gas Cooling after
Sulphur Burner Increased Bed 3 Temperature
Feed Back Loop to Line 40-100HC01-
SW-S1 flow Controller Increase Cooling Water Flow
Rate
Line Manager -
Action By Line
Technician
Reduced Bed 3 Conversion
Structural Damage
No
0
Flow Pipe blockage in 200-
100HS01-FG-P4
Loss Of Production Time Central Bypass Line (Bed 3 and line
200-100HS01-FG-P4&5 can be
bypassed through central pipeline)
Bypass Line and Replace Blocked
Piping
Line Manager -
Action By Line
Technician
Incurred Operating Loss
Pipe Blockage In Stream Plant-Shutdown
As Well As Process
Line Piping
&
Equipment
Contaminants
& Debris
Contaminants Debris or Dirt
In Upstream Offgas
Buildup Of Debris In Piping Filter At Off Gas Line Inlet Attatched
To High Level Flow Indicator & Alarm
Regular Pipe Maintenance &
Replacement Of Corroded
Piping, Valves & Equipment
Line Manager -
Action By Line
Technician
Product Contamination
Rust From Corroded Piping &
Equipment Upstream/On
Process Line
Down Stream Equipment &
Piping Failure
Process
Line Piping
&
Equipment
Water Vapour Insuffcient Gas Drying Mist Formation In Beds Demister Pads At Bed Inlet & Outlet Ensure Inlet Offgas is dry
Process Engineer &
Control Engineer
High Initial water
concentration in off-gas Corrosion In Pipe Gas Drying Before Line Ensure Air Used To Heat Up Beds
During Start-up is dryConverter Shell Corrosion
47. 46
More Process Line
Piping &
Equipment
Flow Increased Off Gas Feed Flow
Rate
Increased Bed Temperature
Temperature Indicator & Controller on Line
Ensure Feedback Loop Is
Properly Calibrated To Control
OFF-GAS temperature in line
Line Manager
- Action By
Line
Technician
Reduced Bed Conversion
Increased Sulphur Buring In
Burner
Potential Pipe Burstage Feed Back Loop to Line 40-100HC01-SW-S1
flow Controller
200-100HS01-FG-P6 Less Process
Operating
Condition
Temperature Fouling In Catalyst Bed Increased Bed Temperature Temperature Indicator & Controller on Line Bypass process fluid around heat
exchanger 1
Line Manager
- Action By
Line
Technician
Excess Gas Cooling after
Sulphur Burner Reduced Bed Conversion
Denatured Catalyst Packing
Temperature Indicator & Controller on Line Ensure Catalyst bed is not heated
above 900K
More Process
Operating
Condition
Temperature Insufficient Gas Cooling after
Sulphur Burner Increased Bed 4 Temperature
Feed Back Loop to Line 40-100HC01-SW-S1
flow Controller Increase Cooling Water Flow
Rate
Line Manager
- Action By
Line
Technician
Reduced Bed 4 Conversion
No
Process Line
Piping &
Equipment
Flow Pipe blockage in 200-
100HS01-FG-P5
Loss Of Production Time Central Bypass Line (Bed 3 and line 200-
100HS01-FG-P6&7 can be bypassed through
central pipeline)
Bypass Line and Replace Blocked
Piping
Line Manager
- Control
Room
Incurred Operating Loss
Pipe Blockage In Stream Plant-Shutdown
As Well As Process Line
Piping &
Equipment
Contaminants
& Debris
Contaminants Debris or Dirt
In Upstream Offgas
Buildup Of Debris In Piping Filter At Off Gas Line Inlet Attached To High
Level Flow Indicator & Alarm
Regular Pipe Maintenance &
Replacement Of Corroded Piping,
Valves & Equipment
Line Manager
- Action By
Line
Technician
Product Contamination
Rust From Corroded Piping &
Equipment Upstream/On
Process Line
Down Stream Equipment &
Piping Failure
Process Line
Piping &
Equipment
Water Vapour Insufficient Gas Drying Mist Formation In Beds Demister Pads At Bed Inlet & Outlet Ensure Inlet Offgass is dry Process
Engineer &
Control
Engineer
High Initial water
concentration in off-gas Corrosion In Pipe Gas Drying Before Line Ensure Air Used To Heat Up Beds
During Start-up is dryConverter Shell Corrosion
More Process Line
Piping &
Equipment
Flow Increased Off Gas Feed Flow
Rate
Increased Bed Temperature
Temperature Indicator & Controller on Line
Ensure Feedback Loop Is
Properly Calibrated To Control
OFF-GAS temperature in line
Line Manager
- Action By
Line
Technician
Reduced Bed Conversion
Increased Sulphur Burning In
Burner
Potential Pipe Burstage Feed Back Loop to Line 40-100HC01-SW-S1
flow Controller
200-100HS01-FG-P7 Less Process
Operating
Condition
Temperature Fouling In Catalyst Bed Increased Bed Temperature Temperature Indicator & Controller on Line Bypass process fluid around heat
exchanger 1
Line Manager
- Action By
Line
Technician
Excess Gas Cooling after
Sulphur Burner Reduced Bed Conversion
Denatured Catalyst Packing
Temperature Indicator & Controller on Line Ensure Catalyst is bed is not
heated above 900K
More Process
Operating
Condition
Temperature Insufficient Gas Cooling after
Sulphur Burner
Reduced Efficiency of Interpass
Absorber
Feed Back Loop to Line 40-100HC01-SW-S1
flow Controller Increase Cooling Water Flow
Rate
Line Manager
- Action By
Line
Technician
Fire Hazard - Potential risk of
Ignition Of Mist & Explosion Temperature Indicator on Line
No
Process Line
Piping &
Equipment
Flow Pipe blockage in 200-
100HS01-FG-P6
Loss Of Production Time Central Bypass Line (Bed 3 and line 200-
100HS01-FG-P6&7 can be bypassed through
central pipeline)
Bypass Line and Replace Blocked
Piping
Line Manager
- Control
Room
Incurred Operating Loss
Pipe Blockage In Stream Plant-Shutdown
As Well As Process Line
Piping &
Equipment
Contaminants
& Debris
Contaminants Debris or Dirt
In Upstream Offgas
Build-up Of Debris In Piping Filter At Off Gas Line Inlet Attached To High
Level Flow Indicator & Alarm
Regular Pipe Maintenance &
Replacement Of Corroded Piping,
Valves & Equipment
Line Manager
- Action By
Line
Technician
Product Contamination
Rust From Corroded Piping &
Equipment Upstream/On
Process Line
Down Stream Equipment &
Piping Failure
