Methanol Production Process Design

1. P R O D U C T I O N
O F G R A D E A A
M E T H A N O L
GROUP 24
SHARMAINE IFFAH NABILAH BINTI MOHD
SHAMSUDIN (153080)
NUR AFZA NURINA BINTI ABDULLAH
SANI (152967)
MAH HIN MAN (148703)
ALICIA CHIN YING JIE (153071)
SUPERVISOR: DR. IRVAN DAHLAN

2. PROCESS ANALYSIS
AND
PROCESS FLOW
DIAGRAM
1

3. CONTENT
Product
Process Alternatives
Comparison of Process Alternatives
Process Description of Chosen Alternative
Demand and Supply for Next 10 Years
Plant Capacity
Process Flow Diagram
Innovative Approach
Conclusion
2

4. METHANO
L • Known as methyl alcohol, abbreviated as MeOH (a hydroxyl group linked to
a methyl group)
• The simplest alcohol with the chemical formula of CH3OH.
• It is a light, volatile, colorless, flammable liquid with a distinctive odor similar to
that of ethanol
• Methanol is widely used as the raw materials for the production of various
chemicals.
• The final products of methanol are found in many household products
PRODUCT
3

5. USAGES OF METHANOL
Usages of Methanol Example
Chemical feedstock Formaldehydes, Acetic acid, MTBE, methanol-to-
olefin (MTO), dimethyl ether (DME), solvents, paints,
cleaning products and etc
Transportation fuel MTBE (anti-knock reagent in gasoline)
Fuel cell technology Fuel for methanol-fuel cell (energy dense, more
energy-efficient)
Wastewater denitrification Remove nitrates from water in denitrification
4

6. 𝑃𝑅𝑂𝐶𝐸𝑆𝑆 𝐴𝐿𝑇𝐸𝑅𝑁𝐴𝑇𝐼𝑉𝐸𝑆
Production of Methanol from
Catalytic Partial Oxidation of
Methane
Production of
Methanol from Coal
Gasification
• Catalytic partial
oxidation of
methane occurs
when methane is
oxidised to form
methanol in the
presence of Pd/Cu
catalyst.
• Coal is converted to
syngas via pyrolysis
and gasification,
then the syngas
reacts in a Cu/ZnO
catalysed reactor to
produce methanol.
• CO2 captured from
flue gas and H2
generated from
electrolysis.
• Catalytic
hydrogenation of
carbon dioxide
takes place to form
methanol and water
in the presence of
catalyst such as
Cu/ZnO
• Hydrocarbons from
natural gas undergo
steam reforming to
form syngas, then
react in a Cu/ZnO
catalysed reactor to
produce methanol.
Production of Methanol from
Catalytic Hydrogenation of Carbon
Dioxide Backed up by Carbon
Dioxide Captured from Flue Gas
Production of Methanol by
Hydrogenation of CO/CO2 from
steam Reforming of Natural Gas
5

7. ALTERNATIVE 1: PRODUCTION OF METHANOL FROM CATALYTIC PARTIAL
OXIDATION OF METHANE
• Pd/Cu catalyst is employed to
catalyse the oxidation of
methane to methanol
Overall equation:
Pd/Cu
CO+O2+→CH3OH+CO2
CO+H2O→CO2+ H2
H2+O2 →H2O2
Pd/Cu
CH4+H2O2 →CH3OH+H2O
PURIFICATION
CATALYTIC
REACTION(Pd/Cu)
600˚C, 30-60 bar
CONDENSATION
GAS-LIQUID
SEPARATION
DISTILLATION
WATER
WATER
METHANOL
CRUDE
METHANOL
METHANOL+
UNREACTED GAS
METHANOL+
UNREACTED GAS
METHANE
NATURAL
GAS
CARBON
MONOXIDE +
AIR
UNREACTED GAS
6

8. ALTERNATIVE 2: PRODUCTION OF METHANOL FROM COAL GASIFICATION
• Pyrolysis of coal: Coal heat → CH4 + C2H6 + CO + CO2 + H2
+ H2O + NH3 + H2S + tar + char
• The oxygen fed into the gasifier will further oxidise char to
produce carbon monoxide and carbon dioxide:
• 2C (s) + O2 (g) → 2CO (g)
• C (s) + O2 (g) → CO2 (g)
• Methanol synthesis: 2H2 (g) + CO (g) ↔ CH3OH (g)
CO2 gasification
C(s)+CO2(g) → 2CO(g)
Steam gasification
C(s)+H2O(g) → CO(g) +H2(g)
Methanation reaction
C(s)+2H2(g) →CH4(g)
Oxygen will also further react with volatiles:
2H2(g)+O2 →2H2O(g)
2CO(g)+O2 →2CO2(g)
2CH4(g)+O2 →2CO(g)+4H2
C2H6+2O2 →2CO(g)+3H2(g)
C6H6+3O2 →6CO(g)+3H2(g)
COAL
GASIFICATION
1560℃
DESULPHURIZING
METHANOL
SYNTHESIS
250℃, 80 bar
GAS-LIQUID
SEPARATION
DISTILLATION
ASH
SULPHUR
POWER TO
HYDROGEN
METHANOL
HYDROGEN
CRUDE SYNGAS
UNREACTED
SYNGAS
METHANOL+
UNREACTED SYNGAS
CRUDE
METHANOL
OXYGEN
WATER
7

9. ALTERNATIVE 3: PRODUCTION OF METHANOL FROM CATALYTIC
HYDROGENATION OF CARBON DIOXIDE BACKED UP BY CARBON DIOXIDE
CAPTURED FROM FLUE GAS
• Carbon dioxide and hydrogen are
fed into a Cu/ZnO catalysed reactor
for the hydrogenation of carbon
dioxide to take place.
• Then, the products are cooled
down in a heat exchanger, prior to
further downstream processing in
distillation column(s) to ensure
production of methanol with high
purity
• CO2 + 3H2 ↔ CO + H2O
POWER-TO-
HYDROGEN
METHANOL
SYNTHESIS
250˚C, 60 bar
CARBON
DIOXIDE
CARBON
CAPTURE UNIT
GAS—LIQUID
SEPARATION
DISTILLATION
CARBON
DIOXIDE
METHANOL
WATER
HYDROGEN
CRUDE
METHANOL PURGE
GAS
UNREACTED
SYNGAS
WATER
FLUE
GAS
CRUDE METHANOL+
UNREACTED SYNGAS
8

10. ALTERNATIVE 4: PRODUCTION OF METHANOL BY HYDROGENATION OF
CO/CO2 FROM STEAM REFORMING OF NATURAL GAS
• Methane in the natural gas reacts
with steam in the presence of
catalyst to produce syngas in a
steam reforming unit. .
• At such condition, methanol is
produced in the reactions shown
below:
• CO + H2 ↔ CH3OH
• CO2 + H2 ↔ H2O + CO
Remaining methane and heavier
hydrocarbons will be converted to
syngas
CH4+0.5O2 ↔ CO+2H2
C2H6+O2 ↔2CO+3H2
C3H8+1.5O2 ↔3CO+4H2
C4H10+2O2 ↔4CO+5H2
C5H12+2.5O2 ↔5CO+6H2
C6H14+3O2 →6CO+7H2
Steam-methane
reforming:
CH4+H2O ↔CO+3H2
Water-gas-shift reaction:
CO+H2O ↔CO2+H2
DESULPHURIZING
STEAM
REFORMING
800-900℃, 15-
30 bar
METHANOL
SYNTHESIS
220-270℃,
50-200 bar
GAS-LIQUID
SEPARATION
DISTILLATION
METHANOL
CRUDE
METHANOL
METHANOL+WATER
+UNREACTED
SYNGAS
SYNGAS
METHANE
NATURAL GAS
STEAM
WATER
PURGE
GAS
UNREACTED
SYNGAS
9

11. ASPECT ALTERNATIVE 1 ALTERNATIVE 2 ALTERNATIVE 3 ALTERNATIVE 4
Sustainability
Flexibility/Controll
ability of
Operation
Economic
Potential and
Feasibility
ASPECT ALTERNATIVE 1 ALTERNATIVE 2 ALTERNATIVE 3 ALTERNATIVE 4
Sustainability • Abundant raw
material
• X = 0.71
• Not abundant
raw material
• X = 0.9
• High cost due
to coal
pyrolysis unit
and constant
transport of
coal
• Abundant raw
material
• X = 0.9
• Relatively new
technologies
in CCUS and
hydrogen
generation
• Abundant raw
material
• X = 0.9
• Commonly
used steam
reforming
unit
ASPECT ALTERNATIVE 1 ALTERNATIVE 2 ALTERNATIVE 3 ALTERNATIVE 4
Flexibility/Controll
ability of
Operation
• “Rigid” as
changes in
temperature
may cause
further
oxidation of
methanol to
formaldehyde
.
• 600°C for
catalytic
partial
oxidation
reactor.
• Mature
control
strategies for
pyrolysis and
gasification
unit.
• 1500°C for
coal pyrolysis
and
gasification
unit; 250°C for
for methanol
reactor.
• CCUS and
hydrogen
generation
unit require
further
research for
more mature
control
strategies
• 250°C for
methanol
synthesis
reactor.
• Mature
control
strategies for
reformers and
methanol
synthesis
reactor.
• 800°C for
reformers,
250°C for
methanol
synthesis
reactor.
COMPARISON OF PROCESS ALTERNATIVES
Among four different alternatives proposed, the chosen
alternative to be the process used in methanol production is by
the hydrogenation of CO/CO2 from steam reforming of natural
gas.
ASPECT ALTERNATIVE 1 ALTERNATIVE 2 ALTERNATIVE 3 ALTERNATIVE 4
Economic
Potential and
Feasibility
• Sufficient
supply of
natural gas,
without
depending on
imports from
foreign
countries.
• Production of
coal is
insufficient.
• Needs to rely
on import of
coal.
• Transport of
coal requires
more cost and
time.
• Scarcity of
data to
validate
feasibility of
large scale
CCUS and
hydrogen
generation
using
electrolysis.
• Sufficient
supply of
natural gas,
without
depending on
imports from
foreign
countries.
10

12. PROCESS DESCRIPTION OF CHOSEN ALTERNATIVE
𝐴𝐿𝑇𝐸𝑅𝑁𝐴𝑇𝐼𝑉𝐸 4
• Steam reforming
and autothermal
reforming:
Rh/MgO/Al2O3
• Methanol synthesis:
Cu/ZnO/Al2O3
• High yield with 0.93
conversion
• Purity of more than
99.85%
Autothermal reforming
CH4+0.5O2 ↔ CO+2H2
C2H6+O2 ↔2CO+3H2
C3H8+1.5O2 ↔3CO+4H2
C4H10+2O2 ↔4CO+5H2
C5H12+2.5O2 ↔5CO+6H2
C6H14+3O2 →6CO+7H2
Steam-methane
reforming:
CH4+H2O ↔CO+3H2
Methanol synthesis
CO + H2 ↔ CH3OH
CO2 + 3H2 ↔ CH3OH + H2O
Steam reforming: 800-900˚C, 15-30 bar
Autothermal reforming: 800-1000˚C, 30-50 b
Methanol synthesis: 220-270˚C, 50-200 bar
Water-gas-shift
reaction:
CO+H2O ↔CO2+H2
11

13. The global
production of
methanol would
reach
137 million tons
in 2025 with
CAGR of 5.66%
per year.
DEMAND AND SUPPLY FOR NEXT 10 YEARS
12

14. PLANT CAPACITY
Global market
production
137 million tonnes
Domestic methanol
production
2.33 million tonnes
Demand of methanol
globally
145 million tonnes
Demand of methanol
in Malaysia in 2020
966 thousand tonnes
Projected Consumption and Demand in 2025 Market Gap:
= Demand of Methanol Globally – Global Market Production
=145 million tonnes-137 million tonnes
=8 million tonnes
Planned Capacity based on domestic production
= 30% of Domestic Methanol Production
= 30% × 2.33 million tonnes
= 690 000 tonnes
Planned Capacity based on market gap:
=8.63% of Market Gap
= 8.63% ×8 million tonnes
= 690 000 tonnes
13

15. PROCESS FLOW DIAGRAM
14

16. AUTOTHERMAL REFORMER
• Autothermal reforming, which
is a secondary reforming unit,
is employed to increase the
conversion of natural gas to
syngas.
• This can increase the overall
yield of methanol and reduce
wastage in the form of
unreacted natural gas
components.
R-102
15

17. FLASH SEPARATOR
• Placing a flash separator before
the effluents from the
autothermal reformer are fed
into the methanol synthesis
reactor.
• The water content in the
reactor feed can be greatly
reduced, which in turn
increases the yield of methanol
V-101
16

18. CONCLUSION
This methanol production facility has an estimated
annual capacity of 690,000 tonnes
Alternative 4 is chosen due to its fairly well-rounded nature in
sustainability, flexibility/controllability and economic potential.
Autothermal reformer and flash separator are the innovative
approaches used in the design
17

19. MASS AND ENERGY
BALANCE
18

20. Feedstock and
products
Feedstock:
Natural Gas, Steam, Oxygen
Products:
Methanol
Operating
period
• 330 working days
• 24 hours/ day
GENERAL REMARKS ON OPERATION
Production
capacity
• 690 000 metric tonnes/year
• 87121.21 kg/hr
19

21. CONTENT
MASS BALANCE
ENERGY BALANCE
UTILITIES
WASTE GENERATION
20

22. MASS BALANCE
General mass balance equation in a system :
𝐼𝑛𝑝𝑢𝑡 + 𝐺𝑒𝑛𝑒𝑟𝑎𝑡𝑖𝑜𝑛 − 𝑂𝑢𝑡𝑝𝑢𝑡 − 𝐶𝑜𝑛𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛 = 𝐴𝑐𝑐𝑢𝑚𝑢𝑙𝑎𝑡𝑖𝑜𝑛
For continuous process at steady state, the accumulation = 0
𝐼𝑛𝑝𝑢𝑡 + 𝐺𝑒𝑛𝑒𝑟𝑎𝑡𝑖𝑜𝑛 − 𝑂𝑢𝑡𝑝𝑢𝑡 − 𝐶𝑜𝑛𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛 = 0
For non-reactive process unit, the generation and consumption = 0
𝐼𝑛𝑝𝑢𝑡 = 𝑂𝑢𝑡𝑝𝑢𝑡
For reactive processes, it is not appropriate to use mass balance for analysis since the atomic balance
may not be the same before and after the unit operation. Therefore, mole balance is used instead:
𝑚𝑜𝑙𝑖𝑛 − 𝑚𝑜𝑙𝑜𝑢𝑡 + 𝑚𝑜𝑙𝑔𝑒𝑛𝑒𝑟𝑎𝑡𝑒𝑑 − 𝑚𝑜𝑙𝑐𝑜𝑛𝑠𝑢𝑚𝑒𝑑 = 0
21

23. MASS BALANCE
Reactive
Equipment
Non-Reactive
Equipment
Reactor
• Steam Reformer, R-101
• Autothermal Reformer, R-102
• Methanol Reactor, R-103
𝐼𝑛𝑝𝑢𝑡 + 𝐺𝑒𝑛𝑒𝑟𝑎𝑡𝑖𝑜𝑛
= 𝑂𝑢𝑡𝑝𝑢𝑡 + 𝐶𝑜𝑛𝑠𝑢𝑚𝑝𝑡𝑖𝑜𝑛
Flash Columns (V-101,V-102)
Reflux Drum, V-103
Distillation Column, T-101
Condenser, E-105
Reboiler, E-106
Heater (E-102,E-104)
Cooler (E-101, E-103)
Compressor, K-101
Pressure Reducing Valve, PIC-101
𝐼𝑛𝑝𝑢𝑡 = 𝑂𝑢𝑡𝑝𝑢𝑡
22

24. MASS BALANCE ANALYSIS FOR REACTIVE SYSTEM
Component Inlet Molar Flow
Rate
Change in Molar Flow Outlet Molar Flow Rate
Limiting Reactant FLR,i −X FLR,f = FLR,i(1 − X)
Reactant FR,i
−θRX FR,f = FR,i(1 − θRX)
Product FP,i +θPX FP,f = FP,i + θPX
Non-involving component FI 0 FI
23

25. MASS BALANCE
Operating Temperature: 300 ℃
Operating Pressure: 50 bar
Methanol Reactor, R-103
24

26. Methanol is produced when the syngas component, CO and 𝐻2 reacts with one another in the presence of
Cu/ZnO/𝐴𝑙2𝑂3 catalyst.
Reaction 1: Methanol Synthesis
𝐶𝑂 + 2𝐻2 ↔ 𝐶𝐻3𝑂𝐻
Reaction 2: Reverse Water- Gas Shift Reaction
𝐶𝑂2 + 𝐻2 ↔ 𝐻2𝑂 + 𝐶𝑂
Methanol Reactor, R-103
Parameter Value
Reaction 1 CO Conversion 0.66818 [5,6]
Reaction 2 𝐶𝑂2 Conversion 0.03353 [5,6]
Selectivity 6.93484
Yield 0.66818
25

27. A simulation was run on MATLAB to verify these values taken for our calculation:
Methanol Reactor, R-103
• Based on MATLAB simulation,
the CO and 𝐶𝑂2 conversion
were found to be 0.700 and
0.00 respectively.
• It is hard to achieve 0%
conversion of reactant from
the side reaction.
• Therefore, the conversion
values used in the manual
calculation of mass balance
analysis were valid.
26

28. A sensitivity analysis is done to evaluate the economic potential by varying the CO conversion in R-103.
Methanol Reactor, R-103
• It is found that the profit
starts to plateau off beyond X
= 0.6.
• It is duly justified to take the
conversion value of about
0.67 for the mass balance
calculation.
27
RM0.00
RM50,000,000.00
RM100,000,000.00
RM150,000,000.00
RM200,000,000.00
RM250,000,000.00
RM300,000,000.00
RM350,000,000.00
RM400,000,000.00
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1
Profit
(MYR)
CO conversion in R-103
Profit (RM) vs CO Conversion in R-103

29. MASS BALANCE
In (Stream 13) Out (Stream 14)
Mass Flow (kg/hr) Mass Fraction Mole Flow (kmol/hr) Mole Fraction Mass Flow (kg/hr) Mass Fraction Mole Flow (kmol/hr) Mole Fraction
Methane, CH4
Ethane, C2H6
Propane, C3H8
Isobutane, iC4H10
Butane, C4H10
Isopentane, iC5H12
Pentane, C5H12
Hexane, C6H14
Carbon dioxide, CO2 524778.46182 0.58523 11926.78322 0.09135 507182.64000 0.56561 11526.87818 0.09221
Nitrogen, N2 28802.58953 0.03212 1028.66391 0.00788 28802.58953 0.03212 1028.66391 0.00823
Oxygen, O2
Carbon Monoxide, CO 116213.93187 0.12960 4150.49757 0.03179 49759.44804 0.05549 1777.12314 0.01422
Hydrogen, H2 226904.76338 0.25304 113452.38169 0.86898 215011.83544 0.23978 107505.91772 0.85997
Methanol, CH3OH 88744.94286 0.09897 2773.27946 0.02218
Water, H2O 7198.29075 0.00803 399.90504 0.00320
Total 896699.74661 1.00000 130558.32639 1.00000 896699.74661 1.00000 125011.76746 1.00000
28

30. MASS BALANCE
Distillation Column, T-101
Operating Pressure: 1 bar
Light Key Component: Methanol
Heavy Key Component: Water
29

31. MASS BALANCE
• The purity of grade AA methanol of 99.85 wt% is achieved using distillation column T-101.
• A “shortcut” method known as Fenske-Underwood-Gilliland-Kirkbride (FUGK) method was used in this
analysis.
• Through mass balance analysis, the amount of methanol and water at the top and bottom product based
on the required purity were determined.
• Based on this composition distribution, top and bottom temperature were determined using dew point and
bubble point calculations respectively.
Distillation Column, T-101
Component Recovery
Methanol 0.98182
Water 0.01818
30

32. Vapor Pressure Equation
𝐶1, 𝐶2, 𝐶3, 𝐶4, 𝐶5 is taken from Table 2-8 Vapor Pressure of Inorganic and Organic Liquids.
P is in Pa and Temperature, T is in Kelvin.
Antoine Equation
𝑃𝑣 = exp 𝐶1 +
𝐶2
𝑇
+ 𝐶3 ln 𝑇 + 𝐶4𝑇𝐶5
where 𝐾𝑖 =
𝑃𝑣𝑎𝑝
𝑃
Raoult’s Law
𝛼𝑖 =
𝐾𝑖
𝐾𝐻2𝑂
• Dew Point (Top temperature)
KH2O,new = ∑
yi
αi
• Bubble Point (Bottom temperature)
KH2O,new = ∑xiαi
Distillation column Top temperature Bottom temperature
C-101 64.41 ℃ 92.53 ℃
31

33. • To obtain 𝜃:
1 − q = ∑
αixi,F
αi−θ
q = 1 since the feed consist only liquid
• Minimum Reflux Ratio
Rmin + 1 = ∑
αixi,D
αi−θ
• Actual Reflux Ratio, R
R = 1.5Rmin = 0.5
Underwood Equation
32

34. MASS BALANCE
In Out
Stream Number 21 26 29
Temperature (°C) 80.00000 64.25222 97.27411
Pressure (bar) 1.00000 1.00000 1.00000
Component
Mass Flow
(kg/hr)
Mass
Fraction
Mole Flow
(kmol/hr)
Mole
Fraction
Mass Flow
(kg/hr)
Mass
Fraction
Mole Flow
(kmol/hr)
Mole
Fraction
Mass Flow
(kg/hr)
Mass
Fraction
Mole Flow
(kmol/hr)
Mole
Fraction
Methanol, CH3OH 88744.94286 0.92497 2773.27946 0.87397 87131.21209 0.99850 2722.85038 0.99734 1613.73077 0.18589 50.42909 0.11382
Water, H2O 7198.29075 0.07503 399.90504 0.12603 130.89313 0.00150 7.27184 0.00266 7067.39762 0.81411 392.63320 0.88618
Total 95943.23360 1.00000 3173.18451 1.00000 87262.10522 1.00000 2730.12222 1.00000 8681.12838 1.00000 443.06229 1.00000
Mass In - Mass
Out
0.00000
33

35. MASS BALANCE
• The outlet stream composition was obtained using the FUG calculation.
• However, the inlet stream composition for this condenser was not known, which necessitates a backward
mass balance calculation.
• The reflux ratio and the boil-up ratio must be determined to find the vapor fraction inside the condenser
and the reboiler.
• Reflux ratio expression: 𝑅 =
𝐿
𝐷
• Boil-up ratio: VB=
R+q xHK,F−xHK,D
xHK,B−xHK,F +q−1
Condenser & Reboiler
34

36. CONTENT
MASS BALANCE
ENERGY BALANCE
UTILITIES
WASTE GENERATION
35

37. ENERGY BALANCE
• According to the second law of thermodynamics, the energy can neither be destroyed nor created as
stated by the Principle of Conservation of energy.
• Energy is typically transferred by heat or work in both open and closed system.
• General Energy Balance Equation
𝐼𝑛𝑝𝑢𝑡 − 𝑂𝑢𝑡𝑝𝑢𝑡 = 𝐴𝑐𝑐𝑢𝑚𝑢𝑙𝑎𝑡𝑖𝑜𝑛
Q − W = ∆Ek + ∆Ep
+ ∆H
• Steady-State Open System Energy Balance
Q − W = ∆H
Whereas −W indicates the work is done by the system; +Q indicates heat supplied to the system
Neglect kinetic and
potential energy change
36

38. • The reference state is set at 25°C and 1 bar.
Reference State
1) Change in kinetic energy and change in potential energy is assumed negligible.
2) Shaft work is assumed negligible due to lack of sufficient information.
3) Specific heat capacities are assumed to be independent of pressure and temperature.
4) Energy transfer efficiency for heat and work is 100%.
5) Heat of formation is involved in only reactors.
6) Heat of mixing and heat of solution are negligible since there is no dissolve of solute in solvent occur.
7) Isothermal reactors for R-101 and R-102.
8) Adiabatic reactor for R-103.
9) Adiabatic pressure changers K-101.
10)Adiabatic mixing at all mixing points and splitting pointes, net heat is zero.
Assumption
37

39. • Can be divided into a few components such as sensible heat, heat of reaction, latent heat and so on
∆H = mCp∆T + ∆Hreaction + ∆Hvaporization
Sensible Heat
• The change in enthalpy due to sensible heat at constant pressure is as the following:
∆H =
T1
T2
CpdT
Latent Heat
• Enthalpy change when one component changes from one phase to another at constant temperature and
pressure.
• Heat of Vaporization is computed using:
∆Hv = C1 1 − Tr
C2+C3Tr+C4Tr
2
Whereas Tr is T / Tc.
Enthalpy, ∆𝐇
38

40. Heat of Reaction
• Heat of formation is chosen to calculate the heat of reaction in the three reactors.
∆Hreaction = ∑noutHformation,out − ∑ninHformation,in
Enthalpy, ∆𝐇
• In compressor K-101, the process is considered as isentropic, which is adiabatic (i.e., Q = 0) and reversible
process.
• Isentropic process equation for work done and new temperature:
W = ∆H
T2
T1
=
P2
P1
1−
1
γ
=
P2
P1
R
Cp
Work, W
39

41. In Out
Stream number 13 14
Temperature (°C) 300.00 300.00
Temperature (K) 573.15 573.15
Pressure (bar) 50.00 50.00
Vapor Fraction 1.00 1.00
Component
Mole flow
(kmol/hr)
Liquid phase
sensible heat
(kJ/mol)
Heat of
vaporisation
(kJ/mol)
Vapor phase
sensible heat
(kJ/mol)
Heat of
formation
(kJ/mol) V ΔP (kJ/mol)
Energy flow
(kJ/kmol)
Mole flow
(kmol/hr)
Liquid phase
sensible heat
(kJ/mol)
Heat of
vaporisation
(kJ/mol)
Vapor phase
sensible heat
(kJ/mol)
Heat of
formation
(kJ/mol) V ΔP (kJ/mol)
Energy flow
(kJ/kmol)
Methane, CH4
Ethane, C2H6
Propane, C3H8
Isobutane, iC4H10
Butane, C4H10
Isopentane, iC5H12
Pentane, C5H12
Hexane, C6H14
Carbon dioxide, CO2 11926.78 11.58 -393.50
-
4555109178.05 11526.88 11.58 -393.50
-
4402376367.31
Nitrogen, N2 1028.66 8.12 0.00 8351651.85 1028.66 8.12 0.00 8351651.85
Oxygen, O2
Carbon Monoxide, CO 4150.50 8.17 -110.52 -424793739.29 1777.12 8.17 -110.52 -181884406.24
Hydrogen, H2 113452.38 7.96 0.00 903334207.64 107505.92 7.96 0.00 855987080.69
Methanol, CH3OH 2773.28 2.66 35.65 13.84 -201.20 -413361760.12
Water, H2O 399.91 20.74 -241.83 -88417005.14
Total 130558.33 0.00 0.00 35.83 -504.02 0.00
-
4068217057.85 125011.77 23.39 35.65 49.68 -947.05 0.00
-
4221700806.28
Total Energy Flow Rate (kJ/hr) -4068217057.85 -4221700806.28
Energy Balance (kJ/hr) -153483748.43
Heat Duty (MJ/hr) -153483.75
Power requirement (kW)
ENERGY BALANCE FOR REACTOR: R-103
Assume R-103 is an isothermal reactor in which the temperature of the reactor is maintained by continuous coolant supply to overcome the
heat released by the reaction
40

42. ENERGY BALANCE
Condenser, E-105
Inlet Temperature: 64.41℃
Outlet Temperature: 64.25 ℃
Operating Pressure: 1 bar
Reboiler, E-106
Inlet Temperature: 92.53℃
Outlet Temperature: 97.27 ℃
Operating Pressure: 1 bar
41

43. ENERGY BALANCE FOR CONDENSER: E-105
E-105 is a total condenser, whereas all components from Stream 22 are condensed fully by the cooling supplied by the condenser
In Out
Stream number 22.00 23.00
Temperature (°C) 64.41 64.25
Temperature (K) 337.56 337.40
Pressure (bar) 1.00 1.00
Vapor Fraction 1.00 0.00
Component
Mole flow
(kmol/hr)
Liquid phase
sensible heat
(kJ/mol)
Heat of
vaporisation
(kJ/mol)
Vapor phase
sensible heat
(kJ/mol)
Heat of
formation
(kJ/mol)
V ΔP (kJ/mol)
Energy flow
(kJ/kmol)
Mole flow
(kmol/hr)
Liquid phase
sensible heat
(kJ/mol)
Heat of
vaporisation
(kJ/mol)
Vapor phase
sensible heat
(kJ/mol)
Heat of
formation
(kJ/mol)
V ΔP (kJ/mol)
Energy flow
(kJ/kmol)
Nitrogen, N2
Oxygen, O2
Carbon Monoxide, CO
Hydrogen, H2
Methanol, CH3OH
4085.51 2.66 35.65 0.35
157922189.5
6
4085.51 3.27 13369733.88
Water, H2O 10.91 2.97 32424.80 10.91 2.96 32292.52
Total
4096.42 5.63 35.65 0.35 0.00 0.00
157954614.3
6
4096.42 6.23 0.00 0.00 0.00 0.00 13402026.40
Total Energy Flow Rate
(kJ/hr) 157954614.36 13402026.40
Energy Balance (kJ/hr) -144552587.96
Heat Duty (MJ/hr) -144552.59
Power requirement (kW)
42

44. ENERGY BALANCE FOR REBOILER: E-106
The boil-up stream is fully reboiled while the bottom product stream is fully in liquid phase.
In Out
Stream number 27.00 28.00 29.00
Temperature (°C) 92.53 97.27 97.27
Temperature (K) 365.68 370.42 370.42
Pressure (bar) 0.00 1.00 1.00
Vapor Fraction 0.11 1.00 0.00
Component
Mole flow
(kmol/hr)
Liquid
phase
sensible
heat
(kJ/mol)
Heat of
vaporisati
on
(kJ/mol)
Vapor
phase
sensible
heat
(kJ/mol)
Heat of
formation
(kJ/mol)
V ΔP
(kJ/mol)
Energy
flow
(kJ/kmol)
Mole flow
(kmol/hr)
Liquid
phase
sensible
heat
(kJ/mol)
Heat of
vaporisati
on
(kJ/mol)
Vapor
phase
sensible
heat
(kJ/mol)
Heat of
formation
(kJ/mol)
V ΔP
(kJ/mol)
Energy
flow
(kJ/kmol)
Mole flow
(kmol/hr)
Liquid
phase
sensible
heat
(kJ/mol)
Heat of
vaporisati
on
(kJ/mol)
Vapor
phase
sensible
heat
(kJ/mol)
Heat of
formation
(kJ/mol)
V ΔP
(kJ/mol)
Energy
flow
(kJ/kmol)
Nitrogen, N2
Oxygen, O2
Carbon Monoxide,
CO
Hydrogen, H2
Methanol, CH3OH
60.99 2.66 35.65 1.74
2442270.5
7
10.57 2.66 35.65 1.98 425574.90 50.43 6.23 313990.11
Water, H2O
474.89 5.09
2418009.0
7
82.26 5.09 39.91
3702099.8
1
392.63 5.09
1999174.7
7
Total
535.89 7.75 35.65 1.74 0.00 0.00
4860279.6
4
92.82 7.75 75.56 1.98 0.00 0.00
4127674.7
1
443.06 11.32 0.00 0.00 0.00 0.00
2313164.8
8
Total Energy Flow
Rate (kJ/hr)
4860279.64 6440839.60
Energy Balance
(kJ/hr)
1580559.96
Heat Duty (MJ/hr) 1580.56
Power
requirement (kW)
43

45. UTILITIES
HEATING AND
COOLING UTILITIES
FUEL GAS ELECTRICITY
44

46. Natural gas is used as fuel to provide heat to steam methane reformer (R-101)
).
Fuel Gas
Steam is used as heating agent for heat exchangers E-102, E-104 and E-106.
Heating Utilities
Unit Code Heating Duty (MJ/hr) Fuel gas required
(kmol/hr)
Fuel gas required (ton/
year)
Estimated cost (RM/yr)
Steam reformer R-101 -558711403.34 789.35 106675.44 193,669,263.37
Code of Equipment Type Heating Duty (MJ/hr) Steam required (ton/ year)
E-102 Superheated steam -787646.57 1778976.88
E-104 Superheated steam -103612.40 660053.422
E-106 Superheated steam -1580.55 10387.26554
Total -7064090353.08 2449417.568
Total water used for heating utility Price
kg/ yr m3/yr RM/ yr
2449417568.26 2449417.57 4,898,835.14
Cost for natural gas
used
Approximately RM 5
millions is required for
heating utilities per
year
45

47. • Cooling water utilities for cooler E-101 and E-103, and condenser E-105 as well as the cooling water utility
equivalent for methanol reactor R-103 are tabulated below.
• To solve for cooling utility, we first presume all the outlet stream of cooling water are recycled to a cooling
tower, then being transported out to cooling unit again. The cooling utility is the make-up water required
accounting for losses in cooling tower.
• The inlet flow rate of the cooling tower is assumed as the sum of the flow rate of all cooling utility required.
𝑀𝑎𝑘𝑒 − 𝑢𝑝 𝑤𝑎𝑡𝑒𝑟 = 𝑒𝑣𝑎𝑝𝑜𝑟𝑎𝑡𝑖𝑣𝑒 𝑙𝑜𝑠𝑠 + 𝑑𝑟𝑖𝑓𝑡 𝑙𝑜𝑠𝑠 + 𝑏𝑙𝑜𝑤𝑑𝑜𝑤𝑛
𝑒𝑣𝑎𝑝𝑜𝑟𝑎𝑡𝑖𝑣𝑒 𝑙𝑜𝑠𝑠 = 0.00085𝑊
𝑐(𝑇1 − 𝑇2)
𝐷𝑟𝑖𝑓𝑡 𝑙𝑜𝑠𝑠 = 0.02% 𝑜𝑓 𝑤𝑎𝑡𝑒𝑟 𝑠𝑢𝑝𝑝𝑙𝑦
𝑏𝑙𝑜𝑤𝑑𝑜𝑤𝑛 =
𝑒𝑣𝑎𝑝𝑜𝑟𝑎𝑡𝑖𝑣𝑒 𝑙𝑜𝑠𝑠 − 𝐶𝑂𝐶 − 1 × 𝑑𝑟𝑖𝑓𝑡 𝑙𝑜𝑠𝑠
𝐶𝑂𝐶 − 1
Cooling Utility
Code Temperature (°C) Heating Duty (MJ/hr) Cooling
equivalence
required (kg/ hr)
Tin Tout
E-101 20 45 797210.33 7612618.522
R-103 30 90 150729.75 599720.48
E-103 20 45 915436.33 8741567.104
E-105 30 57.41 144552.589 1258839.262
2007928.997 18212745.37
46

48. Electricity is supplied to the compressor K-101.
Electricity
Code Power required (kW)
Electricity usage (kWh/yr)
Estimated cost
per year
Peak
(14 hrs)
Off-peak
(10 hrs) RM/ yr
K-101 8100.60 37424759.31 26731970.94 18,012,002.02
Cooling tower make-up water calculations:
Cooling Utility
Cost for cooling
utilities is around RM 9
millions per year
Cost for electricity is
around RM 18 millions
per year
47

49. CONTENT
MASS BALANCE
ENERGY BALANCE
UTILITIES
WASTE GENERATION
48

50. Liquid waste
• The water separated
from syngas-water
separator, V-101
• Methanol-water mixture
generated from the
distillation column, T-
101
Wastewater
• 923,325.45 tons per year is released
Treatment Method
Activated Sludge
• methanol concentration
decreases effectively after
they are consumed by
bacteria (Bacillus
Methanolicus and
Methylophilus
methylotrophus) in the sludge
vessel
Slow Sand Filter
• For methanol concentration of
1ppm, it can successfully
provide the microbial activity
to treat the drinking water
Biologically Activated Filter
• for higher concentrations of
methanol from 100 to 1000
ppm, BAF designed with
counter air flow is needed
49

51. Gaseous waste
• Purge gas produced
from the methanol
reactor effluent
separator
• Flue gas which is
produced by the steam
reformer
Purge Gas
• Most of the separated syngas is
recycled back to stream 18 to
improve conversion and the
remaining syngas is purged and
flared off to prevent inert
accumulation and pressure build up
• excess air is supplied to flare system
so that the purge gas undergoes
complete combustion
Flue Gas
• Composition of flue gas:
• Amount of flue gas produced
• Heat recovery can be implemented
as the flue gas is released at 850 ℃
• Flue gas can be discharged into the
atmosphere without harming the
environment.
Stream
Mass flow rate
kg/yr ton/yr
Stream 17 190259747.49 190,259.75
Flue gas component Mass fraction
Carbon dioxide, CO2 0.15
Nitrogen, N2 0.71
Oxygen, O2 0.01
Argon, Ar 0.01
Water, H2O 0.12
Total 1.00
Mass flow rate
kg/yr ton/yr
1940530852.91 1,940,530.85
50

52. CONCLUSION
The overall mass balance is zero as the inlet total mass flow rate is balanced
by the outlet total mass flow rate
The conversion of CO and 𝐶𝑂2 is 0.66818 and 0.03353 respectively in
Methanol Reactor to produce methanol
The heat duty for reactor R-103 is -153483.75 MJ/hr while -144552.58 MJ/hr
and 1580.55 MJ/hr for condenser and reboiler respectively
Gaseous wastes which is flue gas is the highest waste produced at
1,940,530.85
ton/year
Total utility cost per year is RM 225,775,715.
Preliminary economic potential analysis is RM 352,587,481.95
51

53. REACTOR DESIGN
52

54. R-103
METHANOL REACTOR
53

55. To produce methanol via hydrogenation of syngas
GENERAL INFORMATION
Purpose
P = 50 bars ; T = 300 ℃
Operating
condition
co-precipitated Cu/ZnO/Al2O3
Catalyst
Conversion, X = 0.66181
Data Multi tubular fixed
catalytic reactor
54

56. REACTION KINETICS [1]
Reactions:
𝐶𝑂 + 2H2 ↔ 𝐶H3𝑂𝐻 (main)
Rate Equation:
𝑟3 =
𝑘3𝐾𝐶𝑂 𝑃𝐶𝑂𝑃𝐻2
1.5
−
𝑃𝐶𝐻3𝑂𝐻
𝑃𝐻2
0.5𝐾3
1+𝐾𝐶𝑂𝑃𝐶𝑂+𝐾𝐶𝑂2𝑃𝐶𝑂2 𝑃𝐻2
0.5
+
𝐾𝐻2𝑂
𝑃𝐻2
0.5 𝑃𝐻2𝑂
(main)
CO2 + H2 ↔ H2O + CO (side)
𝑟4 =
𝑘4𝐾𝐶𝑂2 𝑃𝐶𝑂2𝑃𝐻2
1.5
−
𝑃𝐶𝐻3𝑂𝐻𝑃𝐻2𝑂
𝑃𝐻2
1.5𝐾4
1+𝐾𝐶𝑂𝑃𝐶𝑂+𝐾𝐶𝑂2𝑃𝐶𝑂2 𝑃𝐻2
0.5
+
𝐾𝐻2𝑂
𝐾𝐻2
0.5 𝑃𝐻2𝑂
(side)
[1]
[1]
55

57. CATALYST BED SIZING: POLYMATH RESULT
56

58. STEP 1: CATALYST BED SIZING
From polymath, the mass of catalyst required for one single tube is approximately 12 kg. Since the number of
tube used is 1381, the total mass of catalyst required is approximately 12 x 1381 kg = 16572.00 kg.
Assuming the catalyst bed is cylindrical,
𝑉𝑏𝑒𝑑 =
𝑊
𝜌𝑐 1 − 𝜙
Whereas 𝜌𝑐 = 1300 kg/m3-catalyst; 𝜙 = 0.45. Therefore,
𝑉𝑏𝑒𝑑 = 23.18 𝑚3
Component Initial flow
rate (kmol/s)
Final flow
rate
(kmol/s)
Conversion
CO 0.000835 0.000248 0.7033
57

59. STEP 2: COOLING REQUIREMENT
Mineral oil supplied by Radco Industries (XCELTHERM®600) which has the maximum bulk fluid operating
temperature of 316°𝐶 will be used to constantly remove heat to maintain the catalyst packed tube
temperature, Tw at 300°C . Generally, temperature approach of 20°C will be implemented. So, the coolant
return temperature will be set at 280°C . The cooling duty required from energy balance is:
𝑸 = −𝟏𝟓𝟎, 𝟕𝟐𝟗. 𝟕𝟒 𝑴𝑱/𝒉𝒓
Fluid properties
Shell Side Tube Side
Fluid Mineral oil Process fluid
Inlet temperature (°C) 30
300
Outlet temperature (°C) 280
Mean temperature (°C) 155
58

60. STEP 2: COOLING REQUIREMENT [2,3]
The process fluid properties are taken as the average values of those of reactants and products in the tube.
The data are taken from ASPEN PLUS.
Mineral oil properties Value
Density, 𝜌 (kg/m3) 769.7968759
Viscosity, 𝜇 (Pa-s) 0.001014167
Thermal conductivity, k (W/m/°C) 0.125171433
Specific heat capacity, Cp (J/kg/°C) 2456.791229
Prandtl number, Prb 19.90547933
Reactant Product Average
Density, ϱ (kg/m3) 7.111478329 7.42745129 7.26946481
Viscosity, µ (Pa-s) 2.03E-05 2.04E-05 2.03857E-05
Thermal conductivity, k (W/m/°C) 0.238914355 0.233826115 0.236370235
Specific heat capacity, Cp (J/kg/°C) 4514.239624 4443.179332 4478.709478
Prandtl number, Prb 0.38626562
59

61. STEP 2: COOLING REQUIREMENT
Log mean temperature difference (LMTD)
∆𝑇𝐿𝑀𝑇𝐷 =
𝑇1 − 𝑡2 − (𝑇2 − 𝑡1)
ln
𝑇1 − 𝑡2
𝑇2 − 𝑡1
∆𝑇𝐿𝑀𝑇𝐷 =
300 − 280 − (300 − 30)
ln
300 − 280
300 − 30
= 96.05°𝐶
Mass flow rate of coolant required:
𝑄 = 𝑚𝐶𝑝∆𝑇
−150729.74
𝑀𝐽
ℎ𝑟
×
106
𝐽
1𝑀𝐽
×
1ℎ𝑟
3600𝑠
= 𝑚 2456.791229
J
kg − °C
30 − 280 °𝐶
𝑚 = 68.17 𝑘𝑔/𝑠
Heat transfer area, S
Take converged U = 341.49 W/m2/°C, the heat
transfer area is obtained as below:
𝑄 = 𝑈𝑆∆𝑇𝐿𝑀𝑇𝐷
−150729.74
𝑀𝐽
ℎ𝑟
×
106
𝐽
1𝑀𝐽
×
1ℎ𝑟
3600𝑠
= 341.49
𝑊
𝑚2°𝐶
𝑆 96.05°𝐶
𝑆 = 1276.45 𝑚2
60

62. STEP 2: COOLING REQUIREMENT
Fixed V/S ratio
𝑉
𝑆
=
𝑁𝜋𝑟2𝐿
2𝑁𝜋𝑟𝐿
=
𝑟
2
=
𝑑
4
𝑑𝑚𝑎𝑥 =
4𝑉
𝑆
𝑑𝑚𝑎𝑥 =
4 23.18𝑚3
1276.45𝑚2
= 0.07263𝑚 = 72.63𝑚𝑚
Number of tube and tube length
𝑁 =
𝑆
𝜋𝑑𝑜𝑙
Parameter Value
Chosen tube 1.5 in Schedule 80
Inner diameter, di (mm) 38.1
Outer diameter, do (mm) 48.26
Thickness (mm) 10.16
Cross section area, Ac,tube (m2) 0.001829214
Tube length (m) 6.1 (as recommended by Coulson and Richardson)
Number of tube 1380.174798 which is 1381
maximum number of tubes for multitubular reactor used in the industries is 4000 tubes. hence 1381 is
acceptable.
61

63. STEP 2: COOLING REQUIREMENT
Tube arrangement
To minimise pressure drop and for ease of cleaning, the square pitch is used.
The tube pitch is calculated as below:
𝑃𝑡 = 1.25𝑑𝑜
𝑃𝑡 = 1.25 0.04826𝑚 = 0.06033𝑚 = 2.375 𝑖𝑛𝑐ℎ𝑒𝑠
62

64. STEP 2: COOLING REQUIREMENT
The tube bundle diameter is calculated as below:
𝐷𝑏 = 𝑑0
𝑁𝑡
𝐾1
1
𝑛1
Since only one tube pass in established in the reactor, constants used are:
K1 = 0.215 and n1 = 2.207
Therefore, the bundle diameter is:
𝐷𝑏 = 0.04826𝑚
1381
0.215
1
2.207
= 2.56𝑚
The tube in the center row is determined by Db/Pt = 2.56/0.06 = 42.5 which is 43 tubes.
63

65. STEP 2: COOLING REQUIREMENT
Based on the plot, an extrapolation equation has been established to correlate the bundle diameter and
the clearance for fixed-tube type
𝐶𝑙𝑒𝑎𝑟𝑎𝑛𝑐𝑒, 𝐶 𝑚𝑚 = 10𝐷𝑏 + 8 = 10 2.56 + 8
𝐶 = 33.64 𝑚𝑚
Therefore, the shell diameter is:
𝐷𝑠 = 𝐷𝑏 + 𝐶 = 2.56 +
33.64
1000
𝐷𝑠 = 2.60𝑚
64

66. STEP 2: COOLING REQUIREMENT
Generally, a baffle cut of 20 to 25 per cent will be the optimum, giving good heat-transfer rates, without
excessive pressure drop.
The optimum spacing will usually be between 0.3 to 0.5 times the shell diameter.
Spacing of 0.3x Shell diameter is used:
Baffle spacing, lb = 0.3Ds = 0.3(2.60) = 0.78m.
As the tube length used for the catalyst packing is 6.10 m, the number of baffle to be installed can be
calculated as:
𝑁𝑏 =
𝐿𝑡
𝑙𝑏
− 1 =
6.10
0.78
− 1
𝑁𝑏 = 6.83 𝑤ℎ𝑖𝑐ℎ 𝑖𝑠 7 𝑏𝑎𝑓𝑓𝑙𝑒𝑠.
65

67. STEP 2: COOLING REQUIREMENT
Tube-side heat transfer coefficient, hi
𝑁𝑢 =
ℎ𝑖𝑑𝑒
𝑘𝑓
Parameter Value Remark
Mass flow rate (kg/s) 249.08 From mass balance
Volumetric flow rate (m3/s) 34.26 𝑄 = 𝑚/𝜌𝑣
Cross sectional area of one tube (m2) 1.83 × 10−3
Total cross sectional area (m2) 2.53
𝐴𝑡 = 𝑁𝑡 ⋅ 𝐴𝑐
,
𝑜𝑛𝑒 𝑡𝑢𝑏𝑒
Fluid velocity (m/s) 13.56 𝑢 = 𝑄/𝐴𝑡
Equivalent diameter, de = di 0.04
Reynold number 1.84E+05 𝑅𝑒 = 𝜌𝑢𝑑/𝜇
Since the Reynold number, Re > 4000, the flow is turbulent flow, the Nusselt number, Nu can be
determined using equation:
𝑁𝑢 = 𝐶𝑅𝑒0.8
Pr0.33
C = 0.021 for gases
𝑁𝑢 = 2.50 × 102
ℎ𝑖 = 1.55 × 103
𝑊
𝑚2°𝐶
66

68. STEP 2: COOLING REQUIREMENT
Shell side heat transfer coefficient, ho
𝑢 = 𝑗ℎ𝑅𝑒𝑃𝑟0.33
𝜇
𝜇𝑤
0.14
=
ℎ𝑜𝑑𝑒
𝑘𝑓
From the chart above, with baffle cut of 25% and Reynold
number of 6.72 × 104
, heat transfer factor, jh is 𝟏. 𝟖 ×
𝟏𝟎−𝟑.
Parameter Value
As (m2) 0.41
𝑑ℎ(m) 0.048
𝑢𝑠 (m/s) 0.86
Re 6.72 × 104
Parameter Value
Tw (°C) 300
𝜇𝑤 (Pa-s) 2.76 × 10−4
Nu 3.89 × 102
ho (W/m2/°C) 1.02 × 103
ℎ𝑜 = 1.02 × 103
𝑊
𝑚2°𝐶
67

69. STEP 2: COOLING REQUIREMENT
Overall heat transfer coefficient, Uo
The overall heat transfer coefficient is computed as blow:
1
𝑈𝑜
=
1
ℎ𝑜
+
1
ℎ𝑜𝑑
+
𝑑𝑜ln(
𝑑𝑜
𝑑𝑖
)
2𝑘𝑤
+
𝑑𝑜
𝑑𝑖
𝑥
1
ℎ𝑖𝑑
+
𝑑𝑜
𝑑𝑖
𝑥
1
ℎ𝑖
Since the tube material used is the stainless steel 316, the thermal conductivity is between 13-17
W/m °𝐶 An average value of 15 W/m°𝐶 is used for calculation.
The mineral oil in the shell side is considered to be heavy
hydrocarbon with fouling factor of 2000 W/m2 °𝐶 while
the process gases in the tube side is considered as
organic vapor with fouling factor of 5000 W/m2°𝐶.
U = 341.49 W/m2°C ( error = 0%)
Since the error is less than 30% as recommended by
Coulson and Richardson, all the computations above are
correct.
68

70. MECHANICAL DESIGN
69

71. MECHANICAL DESIGN
Construction
Material
• Stainless
Steel 304
Design
Temperature
• 310℃
Design
Pressure
• 55 bar
Welded Joint
Efficiency
• Double
Welded
Butt
Wall Design
• 3.02 mm
Head &
Closure
• Ellipsoidal
Head
L/D ratio
• 2.89
Dead Weight
of Vessel
• 398.44 kN
Vessel
Support
• Conical
Skirt
Nozzle
Design
𝐷𝑖,𝑜𝑝𝑡 = 0.133𝑚𝑉
0.4
𝜇0.13
𝐷 = 1.065
𝑊0.408
𝜌0.343
70

72. MECHANICAL DESIGN [2,3]
Design Pressure and Temperature
Design pressure is taken to be 1.1x operating
pressure.
Design temperature is taken to be the operating
temperature of the reactor +10°C
Vessel Mechanical Design
Design pressure 5.5 N/mm2
Design temperature 310 °C
Material of construction
Welded joint efficiency, J
Material Stainless-steel 304
Design stress 105 N/mm2
Tensile strength 510 N/mm2
Type of joint Double-welded butt joint
Joint efficiency 0.85
71

73. MECHANICAL DESIGN
Tube wall design
Corrosion allowance of 2 mm,
Since the tube diameter is less than 1m, the thickness
computed is more than the minimum thickness needed.
t = 3.02 mm
Tube wall thickness,𝑡𝑡 (mm) 3.02
Reactor vessel head design
For ellipsoidal head,
Corrosion allowance of 2 mm,
Since the reactor vessel head thickness is less than the
shell wall thickness, we take the thickness of the head
to be the same as the shell thickness.
t = 10 mm
Shell wall design
Corrosion allowance of 2 mm,
Since the tube diameter is 2.6 m, the thickness is smaller
than the minimum thickness. Thus, new thickness will be
selected.
t = 10 mm
Tube wall thickness,𝑡𝑡 (mm) 3.37 Wall thickness,𝑡𝑡 (mm) 3.37
74

74. MECHANICAL DESIGN
Maximum Allowable Working Pressure
(MAWP)
𝑀𝐴𝑊𝑃 =
2 × 𝑡𝑒𝑛𝑠𝑖𝑙𝑒 𝑠𝑡𝑟𝑒𝑛𝑔ℎ𝑡 𝑇𝑆 × 𝑗𝑜𝑖𝑛𝑡 𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑐𝑦(𝐽) × 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠(𝑡)
𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝐷 × 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠(𝑡)
Parameter Value
MAWP (N/mm2) 13.13
Design Pressure (N/mm2) 5.5
MAWP
>>>
Operating
pressure
Parameter Value
MAWP (N/mm2) 0.685
Design Pressure (N/mm2) 0.11
Parameter Value
MAWP (N/mm2) 0.685
Design Pressure (N/mm2) 0.11
For tube,
For shell,
For vessel head,
75

75. MECHANICAL DESIGN
Reactor dimension and stress analysis
Length of rector (m) 7.51
Diameter of reactor, 𝐷𝑟(m) 2.62
L/D 2.89
Weight of tubes, 𝑊𝑡 (kN) 213.38
Weight of shell, 𝑊
𝑠 (kN) 22.49
Weight of catalyst (kN) 162.57
Total weight (kN) 398.44
Ratio is still within
the optimum L/D
ratio (2—5)
76

76. MECHANICAL DESIGN
Primary stress analysis
Total longitudinal stress, σz
Principal stress analysis: (𝑁/𝑚𝑚2
)
Maximum compressive stress
Longitudinal stress, 𝜎𝐿 7.21 𝑁/𝑚𝑚2
Circumference stress, 𝜎ℎ 14.43 𝑁/𝑚𝑚2
Direct stress, 𝜎𝑤 4.86 𝑁/𝑚𝑚2
Bending stress, 𝜎𝑏 2.05 𝑁/𝑚𝑚2
Total longitudinal stress
𝜎𝑧(𝑢𝑝𝑤𝑖𝑛𝑑) 4.40 𝑁/𝑚𝑚2
𝜎𝑧(𝑑𝑜𝑤𝑛𝑤𝑖𝑛𝑑) 0.30 𝑁/𝑚𝑚2 Critical buckling stress, 𝜎𝑐 76.41 𝑁/𝑚𝑚2
Maximum compressive stress,
𝜎𝑤
6.91
𝑁/𝑚𝑚2
The greatest principal stress difference is still below the
design stress (105 N/mm2), thus the design in valid!
The maximum compressive stress doesn't exceed the
critical buckling stress, which is 76.41N/mm2, thus, the
design is acceptable.
Principle stresses Upwind Downwind
𝜎1 = 𝜎ℎ 14.43 14.43
𝜎2 = 𝜎𝑧 4.40 0.30
𝜎3 = 0.5𝑃𝑑 0.0550 0.0550
𝜎1 − 𝜎2 10.03 14.13
𝜎1 − 𝜎3 14.37 14.37
𝜎2 − 𝜎3 4.34 -4.04
77

77. MECHANICAL DESIGN
Vessel support:
Dead weight in skirt:
𝜎𝑤𝑠 𝑡𝑒𝑠𝑡 =
𝑊𝐻2𝑂 + 𝑊
𝑣
𝜋(𝐷𝑠𝑘𝑖𝑟𝑡 + 𝑡𝑠𝑘𝑖𝑟𝑡)𝑡𝑠𝑘𝑖𝑟𝑡
𝜎𝑤𝑠 𝑜𝑝𝑒𝑟𝑎𝑡𝑖𝑛𝑔 =
𝑊
𝑣
𝜋(𝐷𝑠𝑘𝑖𝑟𝑡 + 𝑡𝑠𝑘𝑖𝑟𝑡)𝑡𝑠𝑘𝑖𝑟𝑡
Skirt base angle: 𝜃𝑠𝑘𝑖𝑟𝑡 = 𝑡𝑎𝑛−1 𝐻𝑠𝑘𝑖𝑟𝑡
1
2
𝐷𝑠𝑘𝑖𝑟𝑡−𝐷𝑠
Know that
Both stresses satisfy the design criteria, thus the suggested
thickness of the skirt is acceptable.
𝜎𝑠 𝑡𝑒𝑛𝑠𝑖𝑙𝑒 < 𝑓𝐽𝑠𝑖𝑛𝜃𝑠𝑘𝑖𝑟𝑡
𝜎𝑠 𝑐𝑜𝑚𝑝𝑟𝑒𝑠𝑠𝑖𝑣𝑒 < 0.125𝐸
𝑡𝑠𝑘𝑖𝑟𝑡
𝐷𝑠𝑘𝑖𝑟𝑡
𝑠𝑖𝑛𝜃𝑠𝑘𝑖𝑟𝑡
𝜎𝑠(𝑡𝑒𝑛𝑠𝑖𝑙𝑒) 1.08 𝑁/𝑚𝑚2
𝜎𝑠(𝑐𝑜𝑚𝑝𝑟𝑒𝑠𝑠𝑖𝑣𝑒) 13.34 𝑁/𝑚𝑚2
𝑓𝐽𝑠𝑖𝑛𝜃𝑠𝑘𝑖𝑟𝑡 140.24 𝑁/𝑚𝑚2
0.125𝐸
𝑡𝑠𝑘𝑖𝑟𝑡
𝐷𝑠𝑘𝑖𝑟𝑡
𝑠𝑖𝑛𝜃𝑠𝑘𝑖𝑟𝑡
97.30 𝑁/𝑚𝑚2
Since the base angle obtained is still within the
range (80-90°) as suggested by Coulson &
Richardson, the design is valid
Type of skirt Conical skirt
Skirt diameter (m) 2.70 𝑚
Skirt height (m) 4.05 𝑚
Dead weight in skirt (test) 7.50 N/mm2
Dead weight in skirt (operating) 4.68 N/mm2
Skirt base angle, 89.25°
78

78. MECHANICAL DESIGN
Base Ring and Anchor Bolt Design
Approximate pitch circle diameter, 𝐷 = 2 𝑚
Circumference of bolt circle, 𝐶𝑏 = 6283.19 𝑚𝑚
Take minimum bolt spacing = 600 𝑚𝑚
Maximum allowable bolt stress, 𝑓𝑏 125 N/mm2
Maximum moment of bottom skirt, 𝑀𝑠 257851.04 𝑁𝑚
Area of bolt, 𝐴𝑏 78.18 mm2
Diameter of bolt, 𝐷𝑏𝑡 9.98 𝑚𝑚
Maximum allowable ring stress, 𝑓
𝑟 140 N/mm2
𝑁𝑏 =
𝐶𝑏
600
= 10.48 ≈ 12 𝑏𝑜𝑙𝑡𝑠 𝑚𝑢𝑙𝑡𝑖𝑝𝑙𝑒 𝑜𝑓 4
From Coulson & Richardson’s Chemical Engineering
Design, the bolt type selected is M24 as it is the closest
standard size bolt larger than 9.98 mm with a root area
of 78.18 mm2
Compressive load, Fb 158306.65 N/mm
Maximum allowable bearing,Fc 92134.46 N/mm2
Minimum width base ring, 𝐿𝑏 13.16 𝑚𝑚
Actual base ring width,𝑊𝑏 131 𝑚𝑚
Actual bearing pressure, 𝑓
𝑐
′
0.703 𝑁/mm2
Base ring thickness. tbr 10 𝑚𝑚
79

79. For viscous flow in steel pipe, (mineral oil)
𝐷𝑖,𝑜𝑝𝑡 = 0.133𝑚𝑉
0.4
𝜇0.13
For gas, (process fluid)
𝐷 = 1.065
𝑊0.408
𝜌0.343
Stream Schedule No. Dopt (mm) Dopt (in)
Feed inlet 80s 787.4 31.0
Product outlet 80s 787.4 31.0
Mineral oil inlet 40 35.052 1.25
Mineral oil outlet 40 35.052 1.25
MECHANICAL DESIGN
Nozzle design (Kent, 1980)
80

80. SPECIFICATION DATA SHEET FOR REACTOR R-103
METHANOL REACTOR
Identification: R-103 Date: 27/12/2023
Item: Methanol Reactor By: MAH HIN MAN
No. of unit: 1
Function: To produce methanol via hydrogenation of syngas
Operation: Continuous
OPERATING DATA
Materials
handled
(kg/hr):
Feed Outlet
Carbon dioxide,
CO2
524778.46 507182.64
Nitrogen, N2 28802.59 28802.59
Carbon
Monoxide, CO
116213.93 49759.45
Hydrogen, H2 226904.76 215011.84
Methanol,
CH3OH
0.00 88744.94
Water, H2O 0.00 7198.29
Temperature
(°C)
300.0 300.0
Pressure (bar) 50
OPERATIONAL DESIGN
Type of reactor:
Catalytic multi-
tubular reactor
Catalyst: Cu/ZnO/Al2O3
Catalyst bed void: 0.45
Catalyst density
(kg/m3
):
1300.00
Catalyst support: Support grid Catalyst diameter (m): 0.01
REACTOR TUBE DESIGN
Number of tubes: 1381
Mass of catalyst per
tube (kg):
12
Tube inner diameter (m): 0.04 Mass of catalyst (kg): 16572.00
Tube length (m): 6.10 Tube thickness (m): 0.01
Tube pressure drop (bar): 0.49 Shell diameter (m): 2.60
UTILITY DESIGN
Cooling duty (MJ/hr): -150729.74 Coolant: Mineral oil
Heat transfer area (m2
): 1276.44
Coolant flow rate
(kg/hr):
245409.11
Overall heat transfer coefficient
(W/m2
/°C)
341.49
MECHANICAL DESIGN
Design pressure (bar): 55
Design temperature
(°C):
310.0
Materials of construction:
Stainless steel
304
Vessel wall thickness
(mm):
10
Type of head: Ellipsoidal head
Corrosion allowance
(mm)
2
Vessel support: Conical skirt
Feed inlet nozzle size (mm) : 800.1
Product outlet nozzle
size (mm):
800.1
Coolant inlet nozzle size (mm) : 35.052
Coolant outlet nozzle
size (mm):
35.052
81

81. AUTOCAD DRAWING (R-103) 82

82. CONCLUSION 82
Catalyst weight
• 16572.00 kg
Tube length
• 6.1 m
Shell diameter
• 2.6 m
Shell wall thickness
• 10 mm

83. MASS TRANSFER
EQUIPMENT
DESIGN
83

84. CONTENT
GENERAL INFORMATION
DESIGN CALCULATION
SPECIFICATION SHEET
AUTOCAD DRAWING
84

85. T-101 DISTILLATION COLUMN
PURPOSE :
85

86. T-101 DISTILLATION COLUMN
DESIGN CRITERIA
Operating
Pressure
Methanol
Recovery
Type of
tray
Methanol
Purity
Type of
column
1 bar 98.18 % 99.85 wt% Tray
Column
Sieve Tray
86

87. 03
PLATE HYDRAULIC DESIGN
T-101 DISTILLATION COLUMN
Operating
Pressure
01
COMPONENT DISTRIBUTION
02
COLUMN SIZING
04
MECHANICAL DESIGN
DESIGN
PARAMETERS
87

88. Step 1: Dew point, Bubble point & Relative volatilities
Raoult’s Law & Dalton’s Law :
-𝐾𝑖 =
𝑃𝑖,𝑠𝑎𝑡
𝑃
=
𝑦𝑖
𝑥𝑖
Dew point : ∑ 𝑥𝑖 = 1 .
Bubble point : ∑ 𝑦𝑖 = 1.
Relative volatility : 𝛼𝑖,𝐻𝐾 =
𝐾𝑖
𝐾𝐻𝐾
SECTION 1 : COMPONENT DISTRIBUTION [2,3]
Identification Of Heavy Key (HK) and Light Key (LK)
LK – METHANOL
HK - WATER
88

89. Step 2 : Reflux ratio Calculation
 Underwood Equation :
1 − 𝑞 = ∑
𝛼𝑖𝑥𝑖𝑓
𝛼𝑖−𝜃
, solve for θ, θ=1.103.
 Minimum Reflux Ratio:
𝑅𝑚𝑖𝑛 + 1 = ∑
𝛼𝑖𝑥𝑖𝐷
𝛼𝑖−𝜃
= 𝑅𝑚𝑖𝑛 = 0.3336.
 Actual Reflux Ratio:
𝑅 = 1.5𝑅𝑚𝑖𝑛 = 0.5005
Step 3 : Boil up Ratio
 VB =
R+q xHK,F−xHK,D
xHK,B−xHK,F +q−1
; VB= 0.2095
89

90. Step 4 : Minimum No. of stages
𝛼𝐿𝐾,𝐻𝐾 = 𝛼𝐿𝐾,𝐻𝐾
𝐷𝛼𝐿𝐾,𝐻𝐾
𝐵𝛼𝐿𝐾,𝐻𝐾
𝐹 = 3.8423
Using Fenske Equation:
𝑁𝑚𝑖𝑛 =
ln
𝑥𝐿𝐾,𝐷
𝑥𝐻𝐾,𝐷
𝑥𝐻𝐾,𝐵
𝑥𝐿𝐾,𝐵
ln 𝛼𝐿𝐾,𝐻𝐾
= 7.6805 𝑠𝑡𝑎𝑔𝑒𝑠.
Step 5 : Theoretical No. of stages
Using Gilliland Correlation method:
𝑌 =
𝑁 − 𝑁𝑚𝑖𝑛
𝑁 + 1
= 0.75 1 −
𝑅 − 𝑅𝑚𝑖𝑛
𝑅 + 1
0.566
Rearranging the equation:
N =
𝑌+ 𝑁𝑚𝑖𝑛
1−𝑌
= 17.59 ≈ 18 𝑠𝑡𝑎𝑔𝑒𝑠.
90

91. Step 6 : Column Efficiency (O’Connell’s correlation)
𝐸0 = 51 − 32.5 log 𝜇𝐿𝛼𝐿𝐾,𝑎𝑣 = 48.47%
Step 7 : Actual No. of Stages
𝑁𝑎𝑐𝑡 =
𝑁
𝐸0
= 37.13 𝑠𝑡𝑎𝑔𝑒𝑠 ≈ 37 𝑠𝑡𝑎𝑔𝑒𝑠 𝑒𝑥𝑐𝑙𝑢𝑑𝑖𝑛𝑔 𝑟𝑒𝑏𝑜𝑖𝑙𝑒𝑟
Step 8 : Location of Feed Tray
Using Kirkbride equation:
𝑁𝑅
𝑁𝑆
=
𝑥𝐹,𝐻𝐾
𝑥𝐹,𝐿𝐾
𝑥𝐵,𝐿𝐾
𝑥𝐷,𝐻𝐾
2
𝐵
𝐷
0.206
= 2.17
𝑁𝑅 = 2.17 𝑁𝑠
𝑁 = 𝑁𝑅 + 𝑁𝑆 = 37
Solving Simultaneously:
∴ 𝑁𝑅 = 26 𝑠𝑡𝑎𝑔𝑒𝑠 𝑎𝑏𝑜𝑣𝑒 𝑓𝑒𝑒𝑑
𝑁𝑆 = 11 𝑠𝑡𝑎𝑔𝑒𝑠 𝑏𝑒𝑙𝑜𝑤 𝑓𝑒𝑒𝑑
⸫the feed entered at 11th stage from bottom
91

92. Step 1 : Flooding Velocity
Graphical Method from Coulson & Richardson textbook
SECTION 2 : COLUMN SIZING [2,3]
Properties Top Bottom
Vapor Density, 𝜌𝑣 kg/m3 1.138882 0.644467
Liquid density, 𝜌𝐿 kg/m3 748.8824 941.2414
Surface tension, 𝜎 J/m2 0.017391 0.028711
L kmol/h 1366.294 535.8854
V kmol/h 4096.416 92.82314
MW kg/kmol 31.96271 19.59347
FLV 0.013007 0.151066
K1 m/s 0.11 0.09
uf, max m/s 2.740864 3.696122
92

93. Step 2 : Actual Velocity
 85% flooding velocity
Step 3 : Net Column Area
 𝑄 = 𝑢𝑓,𝑛𝑒𝑤𝐴𝑛
Step 4 : Column Cross-sectional Area
Assuming downcomer area as 10%
of net area ,
-𝐴𝑐 =
𝐴𝑛
1−𝑑𝑜𝑤𝑛𝑐𝑜𝑚𝑒𝑟 𝑓𝑟𝑎𝑐𝑡𝑖𝑜𝑛
Step 5 : Column Diameter
 𝐷𝑐 =
4𝐴𝑐
𝜋
uf top (actual) m/s 2.3297
uf bottom (actual) m/s 3.1417
A top m2 13.7076
A bottom m2 13.5081
Qmax top m3/s 31.9350
Qmax bottom m3/s 42.4385
A top (actual) m2 15.2306
A bottom (actual) m2 15.0090
D top m 4.800
D bottom m 4.372
The column diameter is taken to be 4.8m
93

94. SECTION 3 : PLATE HYDRAULIC DESIGN [2,3]
Step 1 : Liquid Flow Pattern
 Liquid Flow Rate = 0.0291
𝑚3
𝑠
, Column Diameter = 4.8 m, From figure 11.28 : Single-pass Cross flow
Step 2 : Provisional Plate Design
 Cross sectional Area, 𝐴𝑐 =
𝜋𝐷2
4
= 20.1062𝑚2
 Downcomer Area, 𝐴𝑑 = 0.1𝐴𝑐 = 2.0106𝑚2
 Active Area, 𝐴𝑎 = 𝐴𝑐 − 2𝐴𝑑 = 16.0850𝑚2
Assuming hole area as 10% of active area
 𝐴ℎ = 0.10 𝐴𝑎 = 1.6085 𝑚2
From figure 11.31 for
𝐴𝑑
𝐴𝑐
= 0.1

𝑙𝑤
𝐷𝑐
= 0.73, ∴ 𝑙𝑤(𝑤𝑒𝑖𝑟 𝑙𝑒𝑛𝑔𝑡ℎ) = 3.504𝑚
Assumption
 Weir height, ℎ𝑤 = 0.05 m
 Hole diameter, 𝑑ℎ = 0.005 m
 Plate thickness, 𝑡𝑝 = 0.005 m
0.73
95

95. Step 3 : Weeping point
 Max liquid flow rate : 27.350 kg/s
 Assuming 70% turndown, min liquid flow rate =
0.7(max flow rate) = 19.145 kg/s
 Max ℎ𝑜𝑤 = 750
𝐿𝑤
𝜌𝑤𝑙𝑤
2/3
= 30.73 𝑚𝑚
 Min ℎ𝑜𝑤 = 750
𝐿𝑤
𝜌𝑤𝑙𝑤
2/3
= 24.22 𝑚𝑚
 ℎ𝑜𝑤 + ℎ𝑤 = 74.22 𝑚𝑚.
 From figure 11.30, 𝐾2 = 30.6
 Vapor velocity through hole = 𝑢ℎ =
𝐾2 −0.90 25.4 − 𝑑ℎ
𝜌𝑣
1/2 = 15.2469
m
s
 Actual minimum vapor velocity = 18.4688 m/s
Since actual minimum vapor velocity >
weeping point, weeping will not occur
30.6
96

96. Step 4 : Plate Pressure Drop
 Vapor velocity through hole,
𝑢ℎ =
𝑄
𝐴ℎ
= 18.4688
𝑚
𝑠
 Plate thickness/hole diameter =1, perforated area
percent = 10, from figure 11.34, Orifice coefficient, 𝐶0 =
0.84.
 Dry plate pressure drop,
ℎ𝑑 = 51
𝑢ℎ
𝐶𝑜
2
𝜌𝑣
𝜌𝐿
= 11.50 𝑚𝑚
 Residual head pressure drop,
ℎ𝑟 =
12.5 × 103
𝜌𝐿
= 13.28 𝑚𝑚
 Total pressure drop,
ℎ𝑡 = ℎ𝑑 + ℎ𝑜𝑤 + ℎ𝑤 + ℎ𝑟 = 105.51 𝑚𝑚
97

97. Step 5 : Downcomer Liquid Back-up
 Head loss in downcomer, ℎ𝑑𝑐 = 166
𝐿𝑤𝑑
𝜌𝐿𝐴𝑚
2
 Take ℎ𝑎𝑝 = ℎ𝑤 − 10 = 40 mm, 𝐴𝑎𝑝 = ℎ𝑎𝑝𝑙𝑤 = 0.1402𝑚2
 𝐴𝑚 = 𝐴𝑎𝑝, 𝑠𝑖𝑛𝑐𝑒 𝐴𝑎𝑝 < 𝐴𝑑 ∴ ℎ𝑑𝑐 = 7.135 𝑚𝑚
 Downcomer liquid back-up, ℎ𝑏 = ℎ𝑜𝑤 + ℎ𝑤 + ℎ𝑡 + ℎ𝑑𝑐 = 0.1942𝑚

𝟏
𝟐
𝒑𝒍𝒂𝒕𝒆 𝒔𝒑𝒂𝒄𝒊𝒏𝒈 + 𝒘𝒆𝒊𝒓 𝒉𝒆𝒊𝒈𝒉𝒕 = 𝟎. 𝟑𝟐𝟓𝒎 > 𝒉𝒃 ∴ 𝒅𝒆𝒔𝒊𝒈𝒏 𝒊𝒔 𝒔𝒂𝒕𝒊𝒔𝒇𝒂𝒄𝒕𝒐𝒓𝒚
Step 6 : Downcomer Residence Time
 𝒕𝒓 =
𝑨𝒅𝝆𝑳𝒉𝒃
𝑳𝒘𝒅
= 𝟐𝟐. 𝟒𝟖𝟖𝟏𝒔 > 𝟑 (𝒔𝒂𝒕𝒊𝒔𝒇𝒂𝒄𝒕𝒐𝒓𝒚)
98

98. Step 7 : Check Entrainment
 Net area vapor velocity,
𝑢𝑛 =
𝑄
𝐴𝑛
= 2.3452 𝑚/𝑠
 Percentage flooding,
𝑢𝑛
𝑣𝑓
× 100% = 74.65%
 From Figure 11.29, when 𝐹𝐿𝑉 = 0.1265, fractional
entrainment, 𝝋 = 0.01𝟑 < 𝟎. 𝟏(Satisfactory)
99

99. Step 8 : No. of Holes
 Area of one hole =
𝜋𝑑ℎ
2
4
= 1.9635 × 10−5
𝑚2
 Number of holes =
𝐴ℎ
𝑎𝑟𝑒𝑎 𝑜𝑓 𝑜𝑛𝑒 ℎ𝑜𝑙𝑒
= 81920 ℎ𝑜𝑙𝑒𝑠
100

100. SECTION 4 : MECHANICAL DESIGN
Construction
Material
• Stainless
Steel 304
Design
Temperature
• 107.3℃
Design
Pressure
• 1.1 bar
Welded Joint
Efficiency
• Double
Welded
Butt
Wall Design
• 12mm
Head &
Closure
• Torispherical
Head
L/D ratio
• 4.747
Dead Weight
of Vessel
• 398.65 kN
Vessel
Support
• Conical Skirt
Nozzle Design
• Following
Kent’s equation
(1980)
101

101. SECTION 4 : MECHANICAL DESIGN [2,3]
Design Pressure and Temperature
Design pressure is taken to be 1.1x operating
pressure.
Design temperature is taken to be the highest
operating temperature of the column, which is
at the bottom, +10°C
Vessel Mechanical Design
Design pressure 0.11 N/mm2
Design temperature 107.3°C
Material of construction
Welded joint efficiency, J
Material Stainless-steel 304
Design stress 165 N/mm2
Tensile strength 510 N/mm2
Type of joint Double-welded butt joint
Joint efficiency 0.85
102

102. SECTION 4 : MECHANICAL DESIGN
Vessel wall design
Corrosion allowance of 2 mm,
For column diameter of 4.8m, this does not suffice,
therefore we take
tt = 12mm
Vessel wall thickness,𝑡𝑡 (mm) 5.883
Vessel head design
For torispherical head,
Corrosion allowance of 2 mm,
Since calculated head thickness is less than column
wall thickness, we take
th = 12mm
Crown radius, 𝑅𝑐 (m) 4.7
Knuckle radius, 𝑅𝐾 (m) 0.282
Stress concentration, 𝐶𝑠 (m) 1.7706
Vessel head thickness, 𝑡ℎ (mm) 5.2615
103

103. SECTION 4 : MECHANICAL DESIGN
Maximum Allowable Working Pressure
(MAWP)
𝑀𝐴𝑊𝑃 =
2 × 𝑡𝑒𝑛𝑠𝑖𝑙𝑒 𝑠𝑡𝑟𝑒𝑛𝑔ℎ𝑡 𝑇𝑆 × 𝑗𝑜𝑖𝑛𝑡 𝑒𝑓𝑓𝑖𝑐𝑖𝑒𝑛𝑐𝑦(𝐽) × 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠(𝑡)
𝑑𝑖𝑎𝑚𝑒𝑡𝑒𝑟 𝐷 × 𝑡ℎ𝑖𝑐𝑘𝑛𝑒𝑠𝑠(𝑡)
Parameter Value
MAWP (N/mm2) 2.16
Design Pressure (N/mm2) 0.11 MAWP
>>>
Operating
pressure
104

104. SECTION 4 : MECHANICAL DESIGN
Column dimension and stress analysis
Total column height, 𝐻𝑇 (m) 22.785
Diameter of column, 𝐷𝑐(m) 4.8
L/D 4.747
Approximate dead weight of
vessel, 𝑊𝑉 (kN)
398.6466
Maximum dead weight of
load, 𝑊𝐿 (kN)
4044.7380
Total weight (testing) (kN) 4443.3846
Ratio is still within
the optimum L/D
ratio (2—5)
105

105. SECTION 4 : MECHANICAL DESIGN
Primary stress analysis
Total longitudinal stress, σz
Principal stress analysis: (𝑁/𝑚𝑚2
)
Maximum compressive stress
Longitudinal stress, 𝜎𝐿 11.0550 𝑁/𝑚𝑚2
Circumference stress, 𝜎ℎ 22.1100 𝑁/𝑚𝑚2
Direct stress, 𝜎𝑤 2.1920 𝑁/𝑚𝑚2
Bending stress, 𝜎𝑏 3.6405 𝑁/𝑚𝑚2
Total longitudinal stress
𝜎𝑧(𝑢𝑝𝑤𝑖𝑛𝑑) 20.0223 𝑁/𝑚𝑚2
𝜎𝑧(𝑑𝑜𝑤𝑛𝑤𝑖𝑛𝑑) -2.3514 𝑁/𝑚𝑚2 Critical buckling stress, 𝜎𝑐 49.7512 𝑁/𝑚𝑚2
Maximum compressive stress,
𝜎𝑤
13.4614
𝑁/𝑚𝑚2
The greatest principal stress difference is still below the
design stress (140.25 N/mm2), thus the design in valid!
The maximum difference doesn't exceed the critical
buckling stress, which is 49.7512 N/mm2, thus, the design
is acceptable.
Principle stresses Upwind Downwind
𝜎1 = 𝜎ℎ 22.1100 22.1100
𝜎2 = 𝜎𝑧 20.0223 -2.3514
𝜎3 = 0.5𝑃𝑑 0.0550 0.0550
𝜎1 − 𝜎2 2.0877 24.4614
𝜎1 − 𝜎3 22.0550 22.0550
𝜎2 − 𝜎3 19.9673 -2.4064
106

106. SECTION 4 : MECHANICAL DESIGN
Vessel support:
Dead weight in skirt:
𝜎𝑤𝑠 𝑡𝑒𝑠𝑡 =
𝑊𝐻2𝑂 + 𝑊
𝑣
𝜋(𝐷𝑠𝑘𝑖𝑟𝑡 + 𝑡𝑠𝑘𝑖𝑟𝑡)𝑡𝑠𝑘𝑖𝑟𝑡
𝜎𝑤𝑠 𝑜𝑝𝑒𝑟𝑎𝑡𝑖𝑛𝑔 =
𝑊
𝑣
𝜋(𝐷𝑠𝑘𝑖𝑟𝑡 + 𝑡𝑠𝑘𝑖𝑟𝑡)𝑡𝑠𝑘𝑖𝑟𝑡
Skirt base angle: 𝜃𝑠𝑘𝑖𝑟𝑡 = 𝑡𝑎𝑛−1 𝐻𝑠𝑘𝑖𝑟𝑡
1
2
𝐷𝑠𝑘𝑖𝑟𝑡−𝐷𝑠
Know that
Both stresses satisfy the design criteria, thus the suggested
thickness of the skirt is acceptable.
𝜎𝑠 𝑡𝑒𝑛𝑠𝑖𝑙𝑒 < 𝑓𝐽𝑠𝑖𝑛𝜃𝑠𝑘𝑖𝑟𝑡
𝜎𝑠 𝑐𝑜𝑚𝑝𝑟𝑒𝑠𝑠𝑖𝑣𝑒 < 0.125𝐸
𝑡𝑠𝑘𝑖𝑟𝑡
𝐷𝑠𝑘𝑖𝑟𝑡
𝑠𝑖𝑛𝜃𝑠𝑘𝑖𝑟𝑡
𝜎𝑠(𝑡𝑒𝑛𝑠𝑖𝑙𝑒) 13.2256 𝑁/𝑚𝑚2
𝜎𝑠(𝑐𝑜𝑚𝑝𝑟𝑒𝑠𝑠𝑖𝑣𝑒) 49.6185 𝑁/𝑚𝑚2
𝑓𝐽𝑠𝑖𝑛𝜃𝑠𝑘𝑖𝑟𝑡 140.2485 𝑁/𝑚𝑚2
0.125𝐸
𝑡𝑠𝑘𝑖𝑟𝑡
𝐷𝑠𝑘𝑖𝑟𝑡
𝑠𝑖𝑛𝜃𝑠𝑘𝑖𝑟𝑡
51.5458 𝑁/𝑚𝑚2
Since the base angle obtained is still within the
range (80-90°) as suggested by Coulson &
Richardson, the design is valid
Type of skirt Conical skirt
Skirt diameter (m) 4.85 𝑚
Skirt height (m) 5.335 𝑚
Dead weight in skirt (test) 33.1278 N/mm2
Dead weight in skirt (operating) 3.2651 N/mm2
Skirt base angle, 89.73°
107

107. SECTION 4 : MECHANICAL DESIGN
Base Ring and Anchor Bolt Design
Approximate pitch circle diameter, 𝐷 = 4.90 𝑚
Circumference of bolt circle, 𝐶𝑏 = 15393. 80 𝑚𝑚
Take minimum bolt spacing = 600 𝑚𝑚
Maximum allowable bolt stress, 𝑓𝑏 125 N/mm2
Maximum moment of bottom skirt, 𝑀𝑠 2441281.76 𝑁𝑚
Area of bolt, 𝐴𝑏 455.4961 mm2
Diameter of bolt, 𝐷𝑏𝑡 24.0823 𝑚𝑚
Maximum allowable ring stress, 𝑓
𝑟 140 N/mm2
𝑁𝑏 =
𝐶𝑏
600
= 25.6563 ≈ 28 𝑏𝑜𝑙𝑡𝑠 𝑚𝑢𝑙𝑡𝑖𝑝𝑙𝑒 𝑜𝑓 4
From Coulson & Richardson’s Chemical Engineering
Design, the bolt type selected is M30 as it is the closest
standard size bolt larger than 24.0823 mm with a root
area of 561mm2
Compressive load, Fb 158306.65 N/mm
Maximum allowable bearing,Fc 7 N/mm2
Minimum width base ring, 𝐿𝑏 22.6152 𝑚𝑚
Actual base ring width,𝑊𝑏 134 𝑚𝑚
Actual bearing pressure, 𝑓
𝑐
′
1.1814𝑁/mm2
Base ring thickness. tbr 12.0923 𝑚𝑚
108

108. SECTION 4 : MECHANICAL DESIGN
Nozzle design (Kent, 1980)
For liquids
𝐷 = 2.607
𝑊
𝜌
0.434
Whereas:
• D is in inches; W is in klbm/hr
Stream W (klbm/hr) r (lbm/ft3) Schedule No. Dopt (mm) Dopt (in)
Feed inlet 211.52 0.068 40s 641.35 25.25
Vapor top outlet 288.66 0.071 40 303.23 11.938
Reflux Inlet 96.28 46.75 40 102.26 4.026
bottom outlet 23.15 0.040 40 128.19 5.047
Boilup inlet 4.01 0.04 40 88.11 3.469
For gases
𝐷 = 1.065
𝑊0.408
𝜌0.343
109

109. AUTOCAD DRAWING OF DISTILLATION COLUMN, T-101
Torispherical head
Plate design
110

110. SPECIFICATION SHEET
DSTILLATION COLUMN
Identification: T-101 Date: 3/1/2014
Item: Distillation Column By: MAH HIN MAN
No. of unit: 1
Function: To purify methanol up to desired purity of 99.85 wt%
Operation: Continuous
OPERATING DATA
Materials
handled:
Feed Distillate Bottom
Methanol
(ton/hr)
88.74 87.13 1.61
Water (ton/hr) 7.20 0.13 7.07
Temperature (°C) 80.0 64.3 97.3
Pressure (bar) 1.01
OPERATIONAL DESIGN
Number of trays: 37 Reflux ratio: 0.5005
Feed point from
bottom:
10 Tray spacing (m): 0.6
Column diameter (m): 4.8 Column height (m): 22.785
Maximum liquid flow
rate (m3
/s):
0.03
Maximum vapor flow rate
(m3
/s):
42.4384577
PLATE HYDRAULIC DESIGN
Active area (m2): 16.0849544 Liquid flow arrangement: Cross flow
Type of tray: Sieve tray Tray thickness (mm): 5
Hole diameter (mm) 5 Weir length (m): 3.504
Active holes: 81920 Weir height (mm): 50
Flow rate turndown
(%):
70
Total plate pressure drop
(mm liquid):
105.512318
Flooding percentage
(%)
85 Entrainment: 0.018
Calming zone width
(mm):
50
Unperforated strip round
plate edge (mm):
50
MCHANICAL DESIGN
Design pressure (bar): 1.1 Design temperature (°C): 107.274114
Materials of
construction
Stainless
steel 304
Column wall thickness (mm): 12
Type of head:
Torisperical
head
Head thickness (mm): 12
Vessel support: Skirt Skirt thickness (mm): 14
Feed inlet nozzle size
(in):
25.25 Corrosion allowance (mm): 2
Reflux inlet nozzle size
(in):
4.026 Top outlet nozzle size (in): 11.938
Boilup inlet nozzle size
(in)
3.469
Bottom outlet nozzle size
(in):
5.047
111

111. CONCLUSION FOR DISTILLATION COLUMN, T-101
• Column length = 22.785m
• Column diameter = 4.8m
• Number of actual stages = 37
• Column efficiency = 48.47%
• No weeping
• No flooding
• Stress analysis passed.

112. REFERENCES
[1] Tonkovich, A.L.Y., Yang, B., Perry, S.T., Fitzgerald, S.P., Wang, Y. (2007). From Seconds to Milliseconds Through
Tailored Microchannel Reactor Design of a Steam Methane Reformer, Catalysis Today, 120 (2007): 21-29.
[2] Sinnott, R. K. (2005). Coulson & Richardson’s Chemical Engineering Design, vol. 6. Elsevier.
[3] Towler, G. P., & Sinnott, R. K. (2008). Chemical Engineering Design: Principles, practice and economics of plant
and Process Design. Elsevier/Butterworth-Heinemann.
[4] Green, D. W., & Perry, R. H. (2008). Perry’s Chemical Engineers’handbook (8th ed.). McGraw-Hill.
[5] Arthur. T (2010). Control structure design for methanol process.
[6] Moulijn, J.A., Makkee, M., van Diepen, A.E. (2014). Chemical Process Technology, 2nd Edition. Doi:
10.1002/cite.201490040.
112