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.