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COAL GAS UREA COMPLEX DESIGN
Kamal MustafaDawood ZafarOsama HasanOun Hasan Syed
Supervisor: Dr. Arshad HussainDepartment of Chemical Engineeirng
SCME, NUST H-12
OBJECTIVETo design a Chemical Production Plant that
produces NH3 and CO2 for urea fertilizer using Coal Syn gas
rather than conventional Natural Gas feedstock.
MOTIVATION
Gas curtailment to fertilizer plants Rising fertilizer demand Hiking fertilizer prices Unexploited coal deposits
PROCESS FLOW DIAGRAM (PFD)
PROCESS DESCRIPTIONSECTION UNIT KEY FOCUS METHOD ASSISTANCE AUXILIARIES
Air Separation
Operation
To separate air into N2
and O2 to ensure supply
into UCG well and Ammonia Reactor
Pressure Swing
Adsorption
Carbon Molecular Sieves
Compressors, De Oxo Reactor,Chilled Water Air
Cooler
Shift Conversion
ProcessConvert CO to CO2 and
maximum hydrogen production
Catalytic Shift Conversion
Iron Oxide Catalyst,Copper Oxide
Catalyst, Steam
Cooling Water, HP Waste heat boiler,BFW Pre Heater
CO2 Removal ProcessTo reduce concentration of CO2 up to 0.05 vol%
on dry basis.Absorption
aMDEA absorbent, Piperazine Activator
Packing Rings
LP Steam, LP BHW Pre-Heater,
Cooling Water
Methanation ProcessTo get synthesis gas free from CO2 and CO up to
less than 5 ppm
Catalytic Bed Reactor
PKR Catalyst(Nickel Based)
Heat Exchangers, Cooler, Separator
Methane Separation
OperationTo separate methane
from hydrogen
Pressure Swing
Adsorption
Carbon Molecular Sieves
Compressors, Chilled Water
Ammonia Synthesis
ProcessTo produce ammonia
from nitrogen and hydrogen
Habers’ Process
Iron Oxide Catalyst
BFW Pre-Heater, Chillers,
Heat Exchangers, Compressors
MATERIAL BALANCEStreams Air PSA inlet
De-oxo reactor inlet
HTSC inlet LTSC inletCO2 Absorber
InletMethanator Inlet
Components Mole% kmol/hr Mole% kmol/hr Mole% kmol/hr Mole% kmol/hr Mole% kmol/hr Mole% kmol/hr
CO - - - - 17.05 71.53 1.70 7.15 0.18 0.72 0.3055 0.7153
CO2 - - - - 23.86 100.14 39.21 164.52 42.25 170.95 0.2190 0.5129
CH4 - - - - 4.54 19.06 4.54 19.06 4.71 19.06 8.14 19.06
H2 - 1.56 8.41 34.09 143.06 49.43 207.44 52.86 213.88 91.33 213.88
O2 21.00 140.11 0.78 4.20 - - - - - - -
H2O - - - - 20.46 85.84 5.11 21.46 0.0033 0.0135 0.0058 0.0135
N2 79.00 527.06 97.66 527.06 - - - - - - - -
NH3 - - - - - - - - - - - -(NH2)2CO - - - - - - - - - - - -
Total 100.00 667.17 100.00 539.67 100.00 419.63 100.00 419.63 100.00 404.62 100.00 234.18
Streams Methane PSA inlet NH3 reactor inlet Urea reactor inlet Urea reactor outlet Recycle Purge
Components
Mole% kmol/hr Mole% kmol/hr Mole% kmol/hr Mole% kmol/hrMole
%kmol/hr Mole% kmol/hr
CO 0.0003 0.0007 - - - - - - - - - -CO2 0.0002 0.0005 - - 33.33 68.75 - - - - - -CH4 8.82 20.29 10.42 144.91 - - - - 13.00 144.30 13.00 0.61H2 91.18 209.69 63.79 886.93 - - - - 61.00 677.24 61.00 2.86O2 - - - - - - - - - - - -
H2O 0.0007 0.0016 - - - - - - - - - -N2 - - 23.39 325.22 - - - - 23.00 255.32 23.00 1.08
NH3 - - 2.40 33.31 66.67 137.50 - - 3.00 33.31 3.00 0.14
(NH2)2CO - - - - - -100.000
069.4444 - - - -
Total 100.00 229.97 100.00 1390.36 100.00 206.25 100.00 69.44100.0
01110.33 100.00 4.69
ENERGY BALANCE
Reaction
Hsupplied
Hremoved
Product
Reactant
Tout
Tin
Trxn
ENERGY BALANCE T in T in a b c d n H n.H*1000 deg C deg C kg mole/hr J/gmol J/hr
CO 25 200 36.11 4.23E-02 -2.9E-05 7.46E-09 7.153 1.21E+02 8.63E+05CO2 25 200 28.95 4.11E-03 3.55E-06 -2.2E-09 164.52 5.16E+03 8.48E+08CH4 25 200 34.31 5.47E-02 3.66E-06 -1.1E-08 19.06 7.09E+03 1.35E+08H2 25 200 28.84 7.65E-05 3.29E-06 -8.7E-10 207.44 5.06E+03 1.05E+09
H2O (g) 25 200 33.46 6.88E-03 7.60E-06 -3.6E-09 21.46 6.01E+03 1.29E+082.34E+04 2162121716
Enthalpy of Syn Gas (kW)600.58936
5
Heating Potential (kW)1159.5901
2 CO + H2O = CO2 + H2
Co-eff 1 1 1 1H f (25 deg C) kJ/gmol -110.52 -241.826 -393.51 0H rxn (25 deg
C) -41.164 kJ/gmol Efficiency 90% H rxn (25 deg C) -37.0476
T in T out a b c d m CpdT m.Cp.dT deg C deg C g mol J/gmol.K kJ/K
CO 25 200 36.11 4.23E-02 -2.9E-05 7.46E-09 1 7.08E+03 7.07877044H2O 25 200 33.46 6.88E-03 7.60E-06 -3.6E-09 1 6.01E+03 6.00975088CO2 25 200 28.95 4.11E-03 3.55E-06 -2.2E-09 1 5.16E+03 5.1557207H2 25 200 28.84 7.65E-05 3.29E-06 -8.7E-10 1 5.06E+03 5.05690913
H Reactants H (CO) + H (H2O) 1.31E+01H Products H (CO2) + H (H2) 1.02E+01
kJ/mol kJ/hr kWH rxn (200 deg C) H p - H r + H rxn(25 deg C) -4.0E+01 -1.1E+06 -317.314
1760.179
ENERGY BALANCE
H in
(kW)
H Sup/Rem
(kW)
H Rxn
(kW)
H Prd. Rem
(kW)
H out
(kW)
HTSC -313.29 161.52 1911.95 - 1760.18
LTSC 1760.18 -1159.59 317.314 - 917.904
LTSC - Absorber 917.904 -775.89 10.6 131.414
Absorber-Methanator 131.414 464.89 57.7 538.604
Methanator 538.604 55.9 594.504
PSA 594.504 -581.125 5.845 7.534
Compressor 7.534 228.02 235.554Compressor - Convertor 238.374 88.73 327.104
Convertor 327.104 675.722 3940.57 4943.396
4202.328
-897.723 6225.734 74.145 9456.194
H out = Hin + H rxn - H pr + H sup/rem 9456.194
EQUIPMENT DESIGN
SHELL AND TUBE HEAT EXCHANGERS
SPECIFICATION SHEET
Heat Duty Q 229520 KJ/hr
Mass of Oil m 4477.15 kg/hr
LMTD LMTD 39.11 °C
Number of Tubes Nt 313 tubes
Tube Side Flow Area at 0.05995 m2
Tube Side HT Co-
Efficienthio 920.94 kJ/hr.m2. °C
Shell side Flow Area as 0.01871 m2
Tube Wall Temperature tw 168.43°C
Shell Side HT Coefficient Ho 1026.6 kJ/hr.m2. °C
Clean Overall
Coefficient,Uc 485.45 kJ/hr.m2.K
Design Overall
Coefficient,UD 469.16 kJ/hr.m2.K
Total Heat Transfer Area A 124.884 m2
Dirt Factor Rd 0.00071
Shell Side Pressure Drop ∆Ps 7.6 psi
Tube Side Pressure Drop ∆PT 0.1203 psi
Shell Side Santotherm
280OC 10OC
Tube Side Syn-Gas
300OC 80OC
SYN GAS COMPRESSORSPECIFICATION SHEET
Average Molecular
WeightMWavg
Compressibility Factor Z
Heat Capacity Ratio k
Inlet Volume V
Volumetric Flowrate Q
Nominal Speed Nnominal
Overall Head H
Outlet Temperature T2
Number of Stages 14
Actual Speed N 1250 rpm
Approximate Power
RequiredG 992.863 kW
ELLIOT / MOLLIER METHOD
TEMPERATURE
80OC 155OC
PRESSURE
4 bar 30 bar
HIGH TEMPERATURE SHIFT CONVERSION
• Iron based catalyst
• Porosity = 0.3
• Conversion 90 %
• Catalyst Weight (Plug Flow Fixed Bed Reactor Equation)
• Rate of Reaction
• Reaction Rate constant
• Equilibrium Constant
SPECIFICATION SHEET
Weight of the catalyst W 714.778 kg
Volume of Catalyst Bed Vc 0.18 m3
Volume of Reactor VR 0.24 m3
Diameter of Reactor D 0.467 m
Height of Reactor H 1.4 m
Catalyst Area Ac 0.17 m2
Height of Catalyst Bed Hc 1.06 m
Volumetric Flow-rate Vo 0.184 m3/s
Space Time τ 1.287 s
Wall Thickness 0.012 m
Density of Gas Mixture ρ11.21
kg/m3
Viscosity of Gas Mixture µ 0.0212 cP
Pressure Drop Across Catalyst
BedΔP 3.3 atm
LOW TEMPERATURE SHIFT CONVERSION
• Copper based catalyst
• Porosity = 0.3
• Conversion = 90%
• Catalyst Weight (Tubular Fixed Bed Reactor Equation)
• Rate of Reaction
• Reaction Rate Constant
• Equilibrium Constant
Weight of the catalyst W 483.53 kg
Volume of Catalyst Bed Vc 0.11 m3
Volume of Reactor VR 0.15 m3
Diameter of Reactor D 0.4 m
Height of Reactor H 1.2 m
Catalyst Area Ac 0.126 m2
Height of Catalyst Bed Hc 0.9 m
Volumetric Flow-rate Vo 0.1538 m3/s
Space Time τ 0,975 s
Wall Thickness 0.01 m
Density of Gas Mixture ρ 13.42 kg/m3
Viscosity of Gas Mixture µ 0.0212 cP
Pressure Drop Across Catalyst
BedΔP 4 atm
ABSORBERLoading Point = 0.5 kmol CO2/kmol MDEA Absorption Factor Ae 1.723
Theoretical Number of Trays N 9
Superficial Velocity VG 0.268 ft/s
Column Efficiency EmV 75%
Actual Number of Trays Nac 12
Height of Column H 6.33 m
Bottom Flooding Velocity Vf 0.333 m/s
Top Flooding Velocity Uaf 0.0625 m/s
Total Diameter of the Column D 1.67 m
Height to Diameter Ratio H/D 3.80
Flow ArrangementCross
FlowSingle Pass
Entrainment ψ 0.0015
Actual Efficiency Ea 74.9%
Total No. of Holes of Sieves 810
Minimum Vapor Velocity 8.46 m/s
Total Plate Pressure Drop Ht
211.45 mm
liq.
Down comer Liquid Backup hb 0.365 m liq.
Residence Time Tr 6.2 s
METHANATOR• Nickel based catalyst
• Porosity = 0.3
• Overall Conversion = 99%
• Catalyst Weight (Tubular Fixed Bed Reactor Equation)
• Rate of Reaction
• Reaction Rate Constant
Weight of the catalyst W 402.76 kg
Volume of Catalyst Bed Vc 0.124 m3
Volume of Reactor VR 0.164 m3
Diameter of Reactor D 0.41 m
Height of Reactor H 1.234 m
Catalyst Area Ac 0.132 m2
Height of Catalyst Bed Hc 0.732 m
Volumetric Flowrate Vo 0.13 m3/s
Space Time τ 1.26 s
Wall Thickness 0.01 m
Density of Gas Mixture ρ 1.44 kg/m3
Viscosity of Gas Mixture µ 0.0201cP
Pressure Drop Across Catalyst
BedΔP 0.32atm
AMMONIA REACTOR• Iron Oxide (Fe2O3) catalyst
• Single Pass Conversion25 %
• Porosity = 0.3
• Catalyst Volume (Fixed Bed RxnEquation)
• Rate of Reaction
Reaction Rate ra
2.109×103
kgmol/m3hr
Weight of the catalyst W 293.7 kg
Volume of Catalyst Bed Vc 0.11 m3
Volume of Bed 1 1 0.02 m3
Volume of Bed 2 2 0.033 m3
Volume of Bed 3 3 0.061 m3
Volume of Reactor VR 0.164 m3
Diameter of Reactor D 0.3 m
Height of Reactor H 2.86 m
Catalyst Area Ac 0.07 m2
Height of Catalyst Bed Hc 1.63 m
Volumetric Flow-rate Vo 0.08 m3/s
Space Time τ 2.53 s
Density of Gas Mixture Ρ 41.8 kg/m3
Viscosity of Gas
Mixtureµ 0.029 cP
Total Pressure Drop ΔP 20.2 atm
INSTRUMENTATION AND PROCESS CONTROL
INSTRUMENTATION AND PROCESS CONTROL
MANIPULATED VARIABLES
Feed stream flow rateTemperature of the reactor
jacketFluid level inside the reactor
DISTURBANCES
Temperature of the reactorPressure of the reactorComposition of streams
CONTROL LOOPS USED
Feed backward control loopFeed forward control loop
HAZOP Study
HAZOP Study
Corrective
Action
Deviation
Possible Cause
Effect
LTSC INLET TEMPERATURE
Deviation High
Possible Cause Low flow rate of Cooling Medium
Effect Catalyst damage
Corrective Action
Increase cooling media flow rate
COST ESTIMATION
COST ESTIMATION
Fixed Capital Cost $
Working Capital
Total Project Cost
Total Variable Cost
Total Fixed Cost $ 4,734,775.94
Direct Production Cost
Annual Production Cost
Production Cost
Total Production Cost
CONCLUSIONS
Process is feasible and is a solution to current Natural Gas Crisis
Pinch Point Analysis – Heat Integration.
Surplus CO2 production for other facilities.
Methane produced can be used as fuel.