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COAL TO LIQUIDPROCESS USINGFISCHER TROPSCHSYNTHESIS
Faheem Shereef
Parnil Singh
Shubam Gupta
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INTRODUCTION
There is a great dearth of liquid fuel in the world
To meet this, the process of coal to liquid was discove
Converting coal to liquid allows coal to be used as an afuel to oil
The scarcity of fuel has led to considerable interest in msynthetic fuels from coalso called coal-to-liquid (CTLlight of coals relatively low prices and the abundance
Moreover, the synthetic fuels provided would be cleancrude oil products displaced
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BENEFITSOF THEPROCESS
Coal is affordable and available worldwide enabling caccess domestic coal reserves and a well-supplied imarket - and decrease reliance on oil imports, imprsecurity.
Coal liquids can be used for transport, cooking, statiogeneration, and in the chemicals industry.
Coal-derived fuels are sulphur-free, low in particulanitrogen oxides.
Liquid fuels from coal provide ultra-cleancooking fuhealth risks from indoor air pollution
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SOMEEXISTINGPROCESSES
Hydrogenation process: In this process, dry coal is mheavy oil recycled from the process.
Pyrolysis and Carbonation process: The carbonizatiooccurs through pyrolysis or destructive distillation.
Direct liquefaction: It works by dissolving the coal in
high temperature and pressure. This process is highly the liquid products require further refining to achieve fuel characteristics.
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THEPROCESS
The process used by us is CTL-RC-CCS: Coal to Liquid
CCS, that is coal is converted to liquid fuel with recycCO2 produced is sequestered.
The main two steps in this process is Coal GasificatioFischer-Tropschprocess
Coal gasification is the process of producing syngasconsisting of methane, carbon-monoxide, hydrogendioxide and water-vapor by reacting coal with a cont
amount of oxygen and steam at high temperatures.
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The reactions involved in the combustion process are:
C+0.5O2COC+ O2CO2
The main gasification reaction occurring is:
C + H2OCO+H2
The Fischer-Tropschprocess is a collection of chemical converts a mixture of carbon-dioxide and hydrogen to liqhydrocarbons. The general reaction can be expressed as
(2n+1)H2 + nCOCnH(2n+2) + nH2O
REACTIONS
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F-T REACTOR
A slurry-phase reactor is used the feed syngas is buban inert oil in which catalyst particles are suspended.
High release rates of reaction heat are accommodateeffective reactor cooling by boiler tubes immersed in t
The vigorous mixing, intimate gas-catalyst contact,
temperature distribution enable a high conversion of fliquids in a relatively small reactor volume.
Conversion by slurry-phase F-T synthesis can achieve fractional conversion of CO to FTL as high as 80%
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PROCESSSTEPS
A coal/water slurry is fed with oxygen from a dedicatcryogenic air separation unit (ASU) into the coal gas
The syngas leaving the reactor is scrubbed with watesome of it enters a single-stage (adiabatic) water gareactor.
CO+H2O CO2+H2
A syngas bypass around the WGS reactor is adjustedthat a molar H2:CO ratio of 1:1 is realized. In this mecaptured CO2 is vented out to the atmosphere.
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For all systems the acid gases CO2, H2S, and COS contsyngas are removed to improve the kinetics and econodownstream synthesis process.
Most of syngas following acid gas removal travels to th
Most of the unconverted syngas and light gases generhydrocarbon recovery step are compressed and recyclthrough the synthesis reactor
The reformed gas stream is mixed with the fresh syng
upstream of the acid gas removal island. A small portiogases leaving the hydrocarbon recovery unit are drawnpurge stream to avoid the build up of inert gases in the
PROCESS STEPS(Cont..)
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COAL-GRADE=ILLINOIS #6
COAL
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DESIGNDESCRIPTION
Equipment
ID
Equipment type Equipment specifications
DECOMP
GASIFIER
T = 25oC, P = 1 atm
GASIFIER Pressure maintained at 30 atm,
ASHREMO
V
CYCLONE (ASH
REMOVAL
SECTION)
Absolute solid separation (solid split f
REFORME
R
STOIC. REACTOR
(WGS SECTION)
Tincalculated from Gasifier outlet, Touisothermally)
COOL1
2-SHIFT WATER
GAS SHIFTREACTION
SECTION
T = 423oC, P = 30 atm
WGS1 T = 423o
C, P = 30 atm; 0o
C temp. appr
COOL2 T = 230oC, P = 30 atm
WGS2 T = 230oC, P = 30 atm; 0oC temp. appr
COND1 CONDENSER T = 37oC, P = 30 atm; water decanted
DESULFUR DESULFURIZER
SEPARATOR
Split fractions of 1 for H2S and NH3rem
PSA MEA ADSORBER S lit f ti f f CO d
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PSA MEA-ADSORBER
(PRESSURE
SWING)
Split fractions of 1 for CO2and rem
mixture
P1 PUMP Pin= 1 atm, Pout= 30 atm; type: cen
efficiency = 0.95
COMP1 COMPRESSOR P = 25 atm;
polytropic efficiency = 0.9, adiabatXE1 HEAT
EXCHANGER
Shell-and-tube heat exchanger wit
FT-REACT FISCHER-
TROPSCH
REACTOR
CSTR reactor to simulate slurry con
loading of 4 kg in vessel volume of
atm; Used Cu-Fe-K catalyst
COOL4 COOLER T = 40oC, P = 30 bar
FLASH FLASH DRUM P = 24 bar; Isothermal flash at 40oC
COMP2 COMPRESSOR P = 27 bar
polytropic efficiency = 0.9, adiabat
ATR AUTO-THERMAL
REFORMER
P = 30 bar, T = 900oC; Gibbs Reacto
COOL3 COOLER T = 40oC, P = 28 bar
DESIGNDESCRIPTION
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PROCESSLIMITATIONS
The product quality is very sensitive to the coal qualityand low carbon content gave very poor outputs
H2:CO ratio should be equal to 1:1 in syngas entering tsynthesis reactor. However this is not ensured and leavariations later
There are some simulation limitations. Fluidization in FT synthesis are not properly accounted for and simul
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PROCESSASSUMPTIONS
No reactions of coke formation are assumed
Heat Capacity is a function of temperature only & not p
100% heat exchange across any units, reactors or heat
No leakage in the system, therefore law of mass conseracross all units.
+=0 &
Heat of reaction is not considered in the energy balanc
100% CO2 absorbed in the absorption tower
Only iso-octane production is taken care of in FT reactoavailability of data
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ENERGYBALANCEFOR GASIFIER
Mass Flow kg/hr Cp
1000.0 300.0kj/kg.k kj/kg.k Kj/h
H2 44.0 0.0 14.42
O2 117.0 300.0 0.92 0.915938
N2 13.0 0.0 1.04
H2O 105.0 0.0 4.19CO 0.0 0.0 0
CO2 0.0 0.0 0
H2S 0.0 0.0 0
H3N 0.0 0.0 0
Total
Mass Flow kg/hr Cp E
1300.0kj/kg.k K
H2 0.0 0
O2 0.0 0
N2 trace 0
H2O 492.0 4.997222
CO 497.8 1.235
CO2 163.5 1.298636
H2S 11.7 1.459706
H3N 13.0 3.708824
Delta H = 4050192 kJ/h
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MASSBALANCEFOR WGS
SECTION
WGS1 WGS2 TH
Stream WG1IN WG1OUT WG2IN WG2OUT IN
Temperature C 972.0 623.0 423.0 230.0 10
Pressure bar 30.4 30.4 30.4 30.4 Mass Flow
kg/hr 2300.0 2300.0 2300.0 2300.0 23
H2 0.3 10.4 10.4 98.3
O2 trace trace trace trace trac
H2O 482.8 311.0 311.0 213.2 15
CO 404.4 386.0 386.0 249.0 5
CO2 163.5 192.4 192.4 407.7 2
H2S 11.7 11.7 11.7 11.7
H3N 13.0 13.0 13.0 13.0
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ENERGYBALANCEFOR WGS
SECTION
Mass Flow kg/hr Cp Inlet
2300.0Kj/kg.k Kj/h
H2 0.3 15.63
O2 trace 0.00
H2O 482.8 5.36
CO 404.4 1.23CO2 163.5 1.29
H2S 11.7 1.44
H3N 13.0 3.65
Mass Flow kg/hr Cp Outl
2300.0Kj/kg.k Kj/h
H2 98.3 14.63
O2 trace 0.00H2O 213.2 4.65
CO 230.6 1.07
CO2 407.7 1.01
H2S 11.7 1.09
H3N 13.0 2.47
Delta H = -2602757.92 kJ/h
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FISCHER-TROPSCHREACTORAND FINAL
SEPARATION Fischer-Tropsch Reactor converts CO and H2to associate
hydrocarbons
Catalyst used was activated Cu-Fe-K/ -Al2O3
Loading of 0.4 kg of catalyst based on catalyst literature
Reference: YN Wang et al, Kinetics modelling of FischerTropsch synthesis over
Cu
K catalyst, 2003
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ENERGYBALANCEFOR FTREACTOR
m (kg/h) Cp (kJ/kgK) Inle
H2 254.40 14.64
O2 trace 0.97
H2O 50.10 4.65
CO 158.90 1.07
CO2 trace 1.02 CH4 0.00 2.96
C2H6 0.00 2.46
I-OCT-01 0.00 2.05
m (kg/h) Cp (kJ/kgK)
Ou
(kJ
H2 4.40 14.65
O2 trace 0.98
H2O 284.40 4.65
CO trace 1.07
CO2 15.70 1.03
CH4 98.80 3.02
C2H6 32.60 2.64
I-OCT-01 35.64 2.05
Delta H = -1146189.05 kJ/h
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AUTO-THERMALREACTOR
Converts light gases (primarily CH4and C2H6) to H2an
high temperatures
Oxygen and Steam input must be optimized for givena fixed temperature
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MASSBALANCEFOR AUTOTHERMAL
REACTOR
ATR-IN ATR-IN THEO OXYGEN OPT STEAM OPT A
Temperature C 40.0 40.0 245.0 185.0
Pressure bar 24.0 24.0 38.0 30.0 Mass Flow kg/hr 435.9 365.6 500.0 220.0
H2 4.4 19.6 0.0 0.0
O2 trace 0.0 500.0 0.0 t
H2O 284.4 212.7 0.0 220.0
CO trace trace 0.0 0.0
CO2 15.7 37.4 0.0 0.0
CH4 98.8 74.6 0.0 0.0
C2H6 32.6 54.4 0.0 0.0
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ENERGYBALANCEFOR AUTOTHERMAL
REFORMER
Mass flow rate Cp
Kg/h Kj/kg.k
H2 4.4 0.0 0.0 14.45 14.64 14.61 19
O2 trace 500.0 0.0 0.00 0.96 0.96
H2O 284.4 0.0 220.0 4.19 4.66 4.42
37
CO trace 0.0 0.0 0.00 0.00 0.00
CO2 15.7 0.0 0.0 0.85 1.01 0.96 4
CH4 98.8 0.0 0.0 2.23 0.00 0.00 690
Mass flow rate Cp Out
kg/h Kj/kg.k KjH2 131.29 15.50
O2 trace 0.00
H2O 114.20 5.14
CO 442.30 1.22
CO2 468.11 1.28
CH4 0.00 0.00Delta E=3295199 kJ/h
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OPTIMIZATIONOF AUTO-THERMALREFORMER
Source: Ekaterini Ch. Vagia, Angeliki A. Lemonidou,Thermodynamic analysis of hydrogen production via
steam reforming
Source: Ekaterini Ch. Vagia,Thermodynamic analysis of hsteam reforming
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HEATEXCHANGENETWORKOPTIMIZATION
Stream
No. Type Ts(oC) Tt(
oC) Cp (Mj/h.K) T (oC
1 Hot 623 423 2.93 -200
2 Hot 423 230 0.30 -193
3 Cold 27.3 245 4.24 217.7
4 Hot 260 73.7 2.20 -186.3
5 Hot 912.7 40 3.94 -872.7
H = (Flow Rate x (Cpt. TtCps. Ts))
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ECONOMICANALYSIS
The sixth-tenths rule is used for ca
costs varying with the capacity
To calculate the cost at a particula
the value at some particular year w
This is done to correct for inflation
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% of equipment
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INDIRECT COST
Sno. Indirect cost% of equipmentcost
1.Engineering andsupervision
2.
Construction expenses
3. Legal expenses
4. Contractors fee
5. Contingency
6. Total indirect costs 1
7. Fixed capital investment 4
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MATERIALSCOST
Cost Cost($)
Operating cost (15% of fixed capital inv
From calculations 1033620
From Aspen 954309
Raw material Cost 39892
Product Cost 184573.2
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CALCULATIONS
Economic Potential = Cost of product-Cost of raw mater
E.P = $144681.2
Economic potential is positive which implies the produeconomically viable
However, Product Cost-Operating Cost= (-$849046.8
Hence our process is infeasible because operating coscovered by the product cost.
The main reason is the large energy requirements of t
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