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Department Of Chemical

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Department Of Chemical

  1. 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. 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. 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. 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. 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. 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. 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
  8. 8. 7 is upstream from the battery region of the heat exchanger and then the heated air at 850o C is passed through each converter bed unit for roughly an hour until there is no temperature fluctuation (steady state). During start up warm air flows through the catalyst beds and down through the line 200-100HS-FG-EM and then the air is vented along with the residual off-gas present in the converter beds to the stack. See Appendix 4 for further information on procedure. Control valves 02-CV-01 to 02-CV-06 are initially closed to pressurise the beds allowing for efficient radial flow of hot air through the beds. Valves 04-CV-01 to 04-CV-05 is partially open to allow hot air to pass through the central core into beds 1-4 and for heat to accumulation in the beds. The pressure is then controlled in each bed by partially opening and closing the control valves at each bed inlet outlet and inlet. This is all done electrically using actuated valves connected to signal transducers, this allows for rapid control response to fluctuations in temperature and pressure. Once the bed reaches its steady state temperature it is necessary to supply off-gas in order to begin normal operation in the converter unit. Care must be taken to not heat the catalyst bed above 900K during the start-up procedure as this will cause thermal degradation of the catalyst in the bed. The furnace is then switched on and allowed to heat up to its normal operating temperature. Then the gas is sent to the reactor via the blower at the beginning of the battery region. ii) Shut-down Procedure During the shut-down procedure the off-gas in the is pumped out of the converter through the central pipeline by pumping in cool dry air through the central core. The pressure build-up causes the off-gas in the converter to flow out of the central core and through of the outlet piping, cooling down the bed. See Appendix 4 for further information on the shutdown procedure. AS A SAFETY FEATURE THE VESSEL CAN BE VENTED IN THE EVENT OF AN EMERGENCY IF RAPID RELEASE OF OFF GAS FROM THE CONVERTER 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 IS USED. 9. References i) J.H. Harker, J.M. Coulson, J.R. Backhurst, J.F. Richardson, (2004) – Coulson and Richardson Volume 6 4th Edition; Sinnott, R.K. © 2005 Elsevier - ii) D K. Louie, (2005) - Handbook Of Sulphuric Acid Manufacture 2nd Edition; DKL Engineering, Inc., 2005 - 0973899204, 9780973899207 iii) W.G Davenport, MJ. King (2006) Sulphuric Acid Manufacture: Analysis, Control And Optimization Volume 13; Elsevier Science & Tech, 2006 - 0080444288, 9780080444284 iv) B Liengme, (2008) - Guide to Microsoft Excel 2007 for Scientists and Engineers 4th Edition; Academic Press, 2008 - 0080923518, 9780080923512 v) R W. Larsen, (2009) Engineering with Excel 3rd Edition; Prentice Hall, 2009 - 0136017754, 9780136017752
  9. 9. 8 Process LineNoPipingMaterialDiameter/mThickness/mmSO2O2CO2N2SO3 P1SS2.02.4653.001.201.67377006936.3085291467.717908840.1184310.073440.0149720.793150212.17 P2SS2.02.4872.021.201.61026774636.3085291490.431029750.0442290.03690.0155620.824430.078871757283.33 P3SS2.02.4688.001.201.61026774636.3085291471.34750580.0442290.03690.0155620.824430.078871757223.54 P4SS2.02.4770.381.201.58629420436.3085291479.890568980.0146720.022350.0157970.836890.110289587250.31 P5SS2.02.4700.001.201.58629420436.3085291472.591939040.0146720.022350.0157970.836890.110289587227.44 P6SS2.02.4723.361.201.57948815736.3085291475.013993260.0061170.018140.0158660.84050.119382867235.03 P7SS2.02.4480.001.201.57948815736.3085291449.777329630.0061170.018140.0158660.84050.119382867155.96 P8SS2.02.4683.001.201.39469560921.5251252841.990214470.0069280.020540.0179680.951860.002704014221.92 P9SS2.02.4702.841.201.39024532621.5251252843.209726590.0005480.01740.0180250.954910.009114825228.36 Process LineNoPipingMaterialDiameter/mThickness/mm S1CS-Sch400.12.4298.004.010.534190839.6154349390.010036988 S2CS-Sch400.13.00298.0036.010.534190839.6154349390.010036988 S3CS-Sch400.23.00416.0036.010.534190839.6154349390.012809107 S4CS-Sch400.23.00416.0036.010.534190839.6154349390.196073306 S5CS-Sch400.23.00355.0036.010.534190839.6154349390.008897981 Bed1Bed2Bed3Bed4 0.670.951.100.85 0.030.030.030.03 0.030.030.030.03 0.020.030.030.03 1.501.501.501.50 GapAboveBedForInletGas&AdditionalFittings/m2.002.002.002.00 4.244.534.684.43 1.001.001.001.00 17.88 ConverterDiameterExcludingThermalInsulation12.08ConverterWallThermalInsulationThickness 10InsulationMaterialMineralWool 16ThermalConductivityW/m/k0.04 FibreBrick ThermalConductivityW/m/k0.98 ReinforcedConcreteHeatTransferCo-efficientofairW/m²/k0.4 SS-ThermalConductivityW/m/K16 CS-ThermalConductivityW/m/K43 WallInsulationThickness/m1.904935195 BrickLayerThickness/m0.205 CentralCoreDiameter/mm NumberOfColumnJoints HeightofSkirtingSupport/mm DepthOfFoundationmm TotalConverterDiametermm TotalConverterHeightmm InternalPExternalPDesignP Bed10.1320.10130.0307627700.91200812.31653604 Bed20.1320.10130.0307627700.91200812.31653604 Bed30.1320.10130.0307627700.91200812.31653604 Bed40.1320.10130.0307627700.91200812.31653604 Total BrickMaterial Insulation Thickness/mm NozzleInlet&OutletDiameters/m TotalConverterHeight/m ProcessPipingSpecification ServicePipingSpecification-ServiceWater ConverterSpecification PackingHeight/m SilicaRockAtSurface/m SilicaRockAtBase/m StainlessSteelGridBase/m 37.18982385 33.45 54.37 61.50 41.73 191.05 24.15 45.08 52.20 32.43 153.86 TotalDeadWeightofBed IncludingValves&Fittings/Mt 9.297455964 9.297455964 9.297455964 9.297455964 Normal Operating Temperature /K PressureSpecsN/mm²DesignTemp ˚C DesignStress N/mm²JointEfficiencyInternalDiameter/mm MechanicalDesignOfConversionVessel SkirtSupportMaterial LoadkN SkirtThicknessmm NumberOfSupportColumns SupportColumnMaterial CarbonSteel SupportColumnDiameter/mm 1874.21 %StreamComposition Normal Operating Temperature /K MolarFlowRate kgmol/s MolarFlowRate kgmol/s MassFlowRatekg/s VolumetricFlow Ratem³/s Normal Operating Pressure /barabs Normal Operating Pressure /barabs MassFlowRatekg/s VolumetricFlow Ratem³/s GapBelowBedForExitGas/m TotalUnitHeight WallThickness /mmShellWeight/Mt PackingWeight /Mt 40 16 500 5000 16 2200 2500 12008 17876.57 FoundationDepth/m NumberOfSupportColums PLANT SPECIFICATION SHEET
  10. 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. 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. 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. 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. 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. 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. 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. 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. 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. 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. 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. 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. 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. 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.
  24. 24. 23 ErgunEquationParameters G D Dp gasρ gasμ 0.6996815 6 0.025 0.529671 0.000055 Bed2Sizing T X(T) K1= K2= K3= Keq= Pa(X(T)) Pb(X(T)) Pc(X(T)) R(T) 1/-R(T) 688 0.00000 67.024494 5.560036 148.0684 345.3526 0.053075 0.044284 0.094646 -0.00039 2541.978 0 0 0 0 700 0.09770 76.8172 4.552022 39.9306 259.1365 0.04789 0.041691 0.099831 -0.00326 306.3989 139.140345 35.67503159 0.751171903 0.035893 710 0.17917 85.759708 3.87305 13.85815 205.4964 0.043566 0.039529 0.104156 -0.01223 81.78646 154.953462 39.72945195 0.8365416 0.039972 720 0.26069 95.450817 3.310171 4.953032 164.0128 0.039239 0.037366 0.108482 -0.02758 36.26429 159.765211 40.96316517 0.862518612 0.041213 730 0.34227 105.92589 2.841296 1.820871 131.7147 0.034909 0.035201 0.112812 -0.03989 25.07124 162.267087 41.60463632 0.8760254 0.041858 740 0.42389 117.22018 2.448924 0.687754 106.4057 0.030577 0.033035 0.117144 -0.04267 23.43439 164.246602 42.11217612 0.886712136 0.042369 750 0.50558 129.36873 2.119118 0.266602 86.45033 0.026241 0.030867 0.12148 -0.03693 27.0817 166.309932 42.64120574 0.89785136 0.042901 760 0.58731 142.4063 1.840722 0.105955 70.62234 0.021904 0.028698 0.125818 -0.02515 39.76006 169.041421 43.34154861 0.912597749 0.043606 770 0.66909 156.36731 1.60476 0.043131 57.99609 0.017563 0.026528 0.130158 -0.00881 113.4809 175.307443 44.94813171 0.946425893 0.045222 771 0.67321 157.81551 1.583205 0.039474 56.88089 0.017344 0.026419 0.130377 -0.00761 131.457 175.812015 45.07750197 0.949149907 0.045352 BED2STOICHIOMETRICTABLE Pressure DropKpa AreaUnder Curve Catalyst Weight/MT CatalystBed Heightm The same method of calculation is carried out for Beds 2,3 and 4 and the results are shown in the tables below: BED 2 0.000 0.005 0.010 0.015 0.020 0.025 0.030 0.035 0.000 0.100 0.200 0.300 0.400 0.500 0.600 0.700 0.800 PressureDrop/Kpa Height / m Bed 1 Pressure Drop Vs Height Profile
  25. 25. 24 0 500 1000 1500 2000 2500 3000 0.00 0.10 0.20 0.30 0.40 0.50 0.60 0.70 0.80 1/R X(T) Bed 2 Levenspiel Plot 0 5 10 15 20 25 30 35 40 45 50 0.00 0.10 0.20 0.30 0.40 0.50 0.60 0.70 0.80 Weight/MetricTonnes Conversion X(T) Bed 2 Catalyst Weight Vs Height
  26. 26. 25 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 0.00 0.10 0.20 0.30 0.40 0.50 0.60 0.70 0.80 Height/m Conversion X(T) Bed 2 Catalyst Height Vs Conversion Bed 3 ErgunEquationParameters G D Dp gas ρ gas μ 0.699681 12.008 0.529671 0.025 0.000055 T X(T) K1= K2= K3= Keq= Pa(X(T)) Pb(X(T)) Pc(X(T)) R(T) 1/-R(T) 700 0 76.8172 4.552022 39.9306 259.1365 0.017606 0.026818 0.132348 -0.00044 2262.3 0 0 0 0 705 0.12492 81.19709 4.196429 23.43558 230.5734 0.015407 0.025718 0.134547 -0.00083 1206.534 216.6633981 18.1534718 0.382238708 0.01826414 710 0.24987 85.75971 3.87305 13.85815 205.4964 0.013207 0.024618 0.136747 -0.00131 764.2713 339.7894622 28.4697761 0.599458358 0.028643336 715 0.37487 90.50948 3.578601 8.255166 183.4419 0.011006 0.023518 0.138948 -0.00162 616.9523 426.1159318 35.7027704 0.75175597 0.035920425 720 0.49991 95.45082 3.310171 4.953032 164.0128 0.008805 0.022417 0.141149 -0.00131 761.1067 512.272176 42.9215022 0.90375327 0.043183164 723.3557 0.58385 98.87682 3.143302 3.529345 152.273 0.007327 0.021678 0.142627 -0.00053 1876.937 622.9908893 52.1982377 1.099083807 0.052516453 Bed2Sizing BED2STOICHIOMETRIC TABLE Pressure Drop Kpa AreaUnder Curve Catalyst Weightkg CatalystBed Heightm
  27. 27. 26 0 500 1000 1500 2000 2500 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 1/R X(T) Bed 3 Levenspiel Plot Bed 3 Levenspiel Plot 0 10 20 30 40 50 60 0 0.2 0.4 0.6 0.8 Weight/MetricTonnes Conversion X(T) BED 3 CATALYST WEIGHT VS CONVERSION BED 3 CATALYST WEIGHT VS CONVERSION
  28. 28. 27 0 0.01 0.02 0.03 0.04 0.05 0.06 0 0.5 1 PressureDrop/Kpa Height/m BED 3 HEIGHT VS PRESSURE DROP BED 3 HEIGHT VS PRESSURE DROP 0 0.2 0.4 0.6 0.8 1 1.2 0 10 20 30 40 50 60 Height/m Weight / Metric Tonnes BED 3 WEIGHT VS HEIGHT BED 3 WEIGHT VS HEIGHT
  29. 29. 28 Bed 4 Ergun Equation Parameters G D gas ρ Dp gas μ 0.699681 12.008 0.529671 0.025 0.000055 T X(T) K1= K2= K3= Keq= Pa(X(T)) Pb(X(T)) Pc(X(T)) R(T) 1/-R(T) 683 0 63.23278 6.05586 259.1276 390.4172 0.008313 0.024646 0.003245 -0.002159581 463.0527 0 0 0 0 685 0.092727 64.72962 5.851558 206.9509 371.6458 0.007542 0.02426 0.004016 -0.002002428 499.3938 44.62239 1.5521 0.040851 0.001952 687 0.1854 66.25288 5.655279 165.4968 353.8784 0.006772 0.023875 0.004786 -0.001893573 528.1022 92.23296 3.208137 0.084438 0.004035 689 0.27812 67.80282 5.466667 132.5182 337.0562 0.006001 0.02349 0.005557 -0.0017993 555.7715 142.4813 4.955926 0.13044 0.006233 691 0.370855 69.37973 5.285383 106.2478 321.1242 0.00523 0.023104 0.006328 -0.00169753 589.091 195.5658 6.802361 0.179038 0.008555 693 0.4636 70.9839 5.111105 85.29398 306.0309 0.004459 0.022719 0.007099 -0.001571241 636.4397 252.3967 8.77911 0.231066 0.011041 695 0.55635 72.61561 4.943527 68.55922 291.7277 0.003688 0.022333 0.00787 -0.001406478 710.9959 314.884 10.95261 0.288272 0.013774 697 0.649121 74.27513 4.782359 55.17694 278.1695 0.002917 0.021948 0.008641 -0.001191534 839.2539 386.7931 13.45382 0.354104 0.01692 699 0.7419 75.96276 4.627323 44.46201 265.3136 0.002146 0.021562 0.009412 -0.000917072 1090.427 476.3101 16.56749 0.436055 0.020836 701 0.83469 77.67877 4.478155 35.87201 253.1202 0.001374 0.021176 0.010184 -0.000576035 1736.006 607.4424 21.12866 0.556105 0.026572 702.8362 0.920189 79.27949 4.346154 29.4866 242.4766 0.000663 0.020821 0.010894 -0.000198896 5027.745 932.3841 32.4311 0.853585 0.040786 703 0.92779 79.42346 4.334604 28.97695 241.5519 0.0006 0.020789 0.010958 -0.000162398 6157.698 974.8943 33.90973 0.892502 0.042646 Bed 2 Sizing BED 2 STOICHIOMETRIC TABLE Area Under Catalyst Weight Catalyst Bed Pressure Drop Kpa 0 1000 2000 3000 4000 5000 6000 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 1/R Conversion X(T) Bed 4 Levenspiel Plot
  30. 30. 29 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 0 5 10 15 20 25 30 35 40 Weight/MetricTonnes Height / m Bed 4 Catalyst Weight Vs Height 0 5 10 15 20 25 30 35 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 Weight/MetricTonnes Conversion X(T) Bed 4 Catalyst Weight Vs Conversion
  31. 31. 30 0 0.005 0.01 0.015 0.02 0.025 0.03 0.035 0.04 0.045 0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1 PressureDropKpa Height / m Bed 4 Height Vs Pressure Drop
  32. 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. 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. 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. 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. 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. 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. 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. 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
  40. 40. 39 Equipment Costing Based On Sizing Calculations - Appendix 3) Source: Seider et al., 2003
  41. 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. 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. 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. 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. 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. 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. 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

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