1. PRODUCTION OF 100MT DISTILLED MONOGLYCERIDE (DMG)
SAJJAD KHUDHUR ABBAS
Ceo , Founder & Head of SHacademy
Chemical Engineering , Al-Muthanna University, Iraq
Oil & Gas Safety and Health Professional – OSHACADEMY
Trainer of Trainers (TOT) - Canadian Center of Human
Development

2. PROCESS BACKGROUND

3. Formed biochemically via release of a fatty acid from diacylglycerol by diacylglycerol lipase.
Monoglyceride (MG) - chemical compound a.k.a monoacylglycerol
Industrial chemical and biological processes.
General Information
Act as emulsifiers - mix ingredients that would not otherwise blend well

5. PROCESS DESCRIPTION

7. Monoglyceride synthesis
Glycerolysis procedure is more economical - fats are cheaper and less glycerol is required.
Fats and fatty acids are insoluble in glycerol - high temperatures are required to force the reaction to proceed.
On production scale, direct esterification and interesterification can be done continuously or batchwise.

8. Flow chart

9. Physical and Chemical Properties

10. DIGLYCERIDES (DISTEARIN) White to pale yellow Wax-like solid
COMPONENTS
Appearance
Formula MW
(g/mol) Tb
(K) Tf
ΔfHo298
(kJ/mol)
GLYCEROL
Clear viscous liquid
Little or no odor
C3H5(OH)3
444
472
MONOGLYCERIDES
(MONOSTEARIN)
Colorless
Odorless
Sweet-taste
Flaky powder
C21H4204
940.09
424.9
DIGLYCERIDES
(DISTEARIN)
White to pale yellow
Wax-like solid
Mild fatty odour
C39H76O5
454.8

11. LEVEL 1 SELECTION OF PROCESSING MODE

12. Proposed Process
Batch
Continuous
Operating 24 hr/day
Production is continuous
Total batch time 3-5 hours
7 batches/day production
99% purity
% purity
98% purity
Annual cost is higher
Annual cost is lower
Lower maintenance cost
Higher specific manufacturing and operating cost
Higher maintenance cost

13. LEVEL 2 SELECTION OF INPUT-OUTPUT STRUCTURE

14. Reaction Information Reaction 1 Reaction 2 Rate constant 350oF 460oFk1
0.291
1.566 k2
0.163
0.220

15. Input-Output Structure

16. Glycerol Selectivity r1=k1CGCFA(1) r2=k2CMCFA(2)
Base on consecutive reaction
Glycerol
− 𝑑𝐶 𝐺 𝑑𝑡 =− 𝑟 𝐺 = 𝑘 1 𝐶 𝐺 𝐶 𝐹𝐴 (3)
Fatty acid
− 𝑑𝐶 𝐹𝐴 𝑑𝑡 =− 𝑟 𝐹𝐴 = 𝑟 1 + 𝑟 2 = 𝑘 1 𝐶 𝐺 𝐶 𝐹𝐴 + 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 (4)
Monoglyceride
𝑑𝐶 𝑀 𝑑𝑡 = 𝑟 𝑀 = 𝑟 1 − 𝑟 2 = 𝑘 1 𝐶 𝐺 𝐶 𝐹𝐴 − 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 (5)
Water
𝑑𝐶 𝑊 𝑑𝑡 = 𝑟 𝑊 = 𝑟 1 + 𝑟 2 = 𝑘 1 𝐶 𝐺 𝐶 𝐹𝐴 + 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 (6)
Diglyceride
𝑑𝐶 𝐷 𝑑𝑡 = 𝑟 𝐷 = 𝑟 2 = 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 (7)

17. 𝐶 𝐺 = 𝐶 𝐺𝑜 (1− 𝑋 𝐺 )(8)
− 𝑑𝐶 𝐺 𝑑𝑡 = 𝑑 (𝐶 𝐺𝑜 (1− 𝑋 𝐺 )) 𝑑𝑡 = 𝐶 𝐺𝑜 𝑑 𝑋 𝐺 𝑑𝑡 = 𝑘 1 𝐶 𝐺𝑜 1− 𝑋 𝐺 𝐶 𝐹𝐴
𝑑 𝑋 𝐺 𝑑𝑡 = 𝑘 1 1− 𝑋 𝐺 𝐶 𝐹𝐴 (9)
− 𝑑𝐶 𝐹𝐴 𝑑𝑡 = 𝑘 1 𝐶 𝐺𝑜 (1− 𝑋 𝐺 ) 𝐶 𝐹𝐴 + 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 (10)
𝑑𝐶 𝑀 𝑑𝑡 = 𝑘 1 𝐶 𝐺𝑜 (1− 𝑋 𝐺 ) 𝐶 𝐹𝐴 − 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 (11)
𝑑𝐶 𝑊 𝑑𝑡 = 𝑘 1 𝐶 𝐺𝑜 (1− 𝑋 𝐺 ) 𝐶 𝐹𝐴 + 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 (12)
𝑑𝐶 𝐷 𝑑𝑡 = 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 (13)

18. By applying chain rule;
− 𝑑𝐶 𝐺 𝑑 𝑋 𝐺 = 𝑘 1 𝐶 𝐺𝑜 1− 𝑋 𝐺 𝐶 𝐹𝐴 𝑘 1 1− 𝑋 𝐺 𝐶 𝐹𝐴 = 𝐶 𝐺𝑜 (14)
− 𝑑𝐶 𝐹𝐴 𝑑 𝑋 𝐺 = 𝑘 1 𝐶 𝐺𝑜 1− 𝑋 𝐺 𝐶 𝐹𝐴 + 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 𝑘 1 1− 𝑋 𝐺 𝐶 𝐹𝐴 = 𝐶 𝐺𝑜 + 𝑘 2 𝐶 𝑀 𝑘 1 1− 𝑋 𝐺 (15) 
𝑑𝐶 𝑀 𝑑 𝑋 𝐺 = 𝑘 1 𝐶 𝐺𝑜 1− 𝑋 𝐺 𝐶 𝐹𝐴 − 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 𝑘 1 1− 𝑋 𝐺 𝐶 𝐹𝐴 = 𝐶 𝐺𝑜 − 𝑘 2 𝐶 𝑀 𝑘 1 1− 𝑋 𝐺 (16)
𝑑𝐶 𝑊 𝑑 𝑋 𝐺 = 𝑘 1 𝐶 𝐺𝑜 1− 𝑋 𝐺 𝐶 𝐹𝐴 + 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 𝑘 1 1− 𝑋 𝐺 𝐶 𝐹𝐴 = 𝐶 𝐺𝑜 + 𝑘 2 𝐶 𝑀 𝑘 1 1− 𝑋 𝐺 (17)
𝑑𝐶 𝐷 𝑑 𝑋 𝐺 = 𝑘 2 𝐶 𝑀 𝐶 𝐹𝐴 𝑘 1 1− 𝑋 𝐺 𝐶 𝐹𝐴 = 𝑘 2 𝐶 𝑀 𝑘 1 1− 𝑋 𝐺 (18)

20. Economic Potential of Level 2

21. LEVEL 3: REACTOR AND RECYCLE STREAMS

22. Generally, there will be input for the process and output from the process. Here we can define what are the related variables or input-output that present in this process.
Feed stream: In this process, the feed raw material is assumed already pure, so no need to purify the feed streams.
Excess reactant: fatty acid is fed as an excess reactant and is supplied in liquid form.
Recycle and purge: There are recycle stream from glycerol and fatty acid but there are no purges from the process.

23. Recycle Stream

25. Adiabatic Temperature
Energy balances
Simplified;
Where;

26. Where from the process,

27. 𝑇 𝑎 = 𝑇 𝑚 − 𝑗−1 𝑁 𝑛 𝑗 ∆𝐻 𝑟𝑗 𝑚 𝑖−1 𝑀 𝑃 𝑖 𝑐 𝑝𝑖
𝑗=1 𝑁 𝑛 𝑗 ∆𝐻 𝑟𝑗 𝑚 = 𝐹 𝐹𝐺 𝑆 𝑀 𝑋 𝐺 ∆𝐻 𝑟1 ° + 𝐹 𝐹𝐺 𝑆 𝑀 𝑋 𝐺 ∆𝐻 𝑟2 ° + 𝐹 𝐹𝐺 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝 𝑀 + 𝐹 𝐹𝐺 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝 𝐷 + 𝐹 𝐹𝐺 𝑋 𝐺 𝑐 𝑝 𝑊 − 𝐹 𝐹𝐺 1− 𝑋 𝐺 𝑐 𝑝 𝐺 − 𝐹 𝐹𝐺 1− 𝑋 𝐺 𝑐 𝑝 𝐹𝐴 ( 𝑇 𝑚 −25)

28. 𝑖=1 𝑀 𝑃 𝑖 𝑐 𝑝𝑖 = 𝐹 𝐹𝐺 1− 𝑋 𝐺 𝑐 𝑝𝐺 + 𝐹 𝐹𝐺 1− 𝑋 𝐺 𝑐 𝑝𝐹𝐴 + 𝐹 𝐹𝐺 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝𝑀 + 𝐹 𝐹𝐺 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝𝐷 + 𝐹 𝐹𝐺 𝑋 𝐺 𝑐 𝑝𝑊 𝑖=1 𝑀 𝑃 𝑖 𝑐 𝑝𝑖 = 𝐹 𝐹𝐺 1− 𝑋 𝐺 𝑐 𝑝𝐺 + 𝐹 𝐹𝐺 1− 𝑋 𝐺 𝑐 𝑝𝐹𝐴 + 𝐹 𝐹𝐺 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝𝑀 + 𝐹 𝐹𝐺 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝𝐷 + 𝐹 𝐹𝐺 𝑋 𝐺 𝑐 𝑝𝑊 𝑇 𝑎 = 𝑇 𝑚 − 𝑆 𝑀 𝑋 𝐺 ∆𝐻 𝑟1 ° + 𝑆 𝑀 𝑋 𝐺 ∆𝐻 𝑟2 ° + 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝 𝑀 + 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝 𝐷 + 𝑋 𝐺 𝑐 𝑝 𝑊 − 1− 𝑋 𝐺 𝑐 𝑝 𝐺 − 1− 𝑋 𝐺 𝑐 𝑝 𝐹𝐴 ( 𝑇 𝑚 −25) 1− 𝑋 𝐺 𝑐 𝑝𝐺 + 1− 𝑋 𝐺 𝑐 𝑝𝐹𝐴 + 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝𝑀 + 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝𝐷 + 𝑋 𝐺 𝑐 𝑝𝑊

29. XG
Ta(K)
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
0.9 1

31. Isothermal heat load can be obtained from
𝑄= 𝑗=1 𝑁 𝑛 𝑗 ∆𝐻 𝑟𝑗 𝑚 + 𝑖=1 𝑀 𝑃 𝑖 𝑐 𝑝𝑖 ( 𝑇 𝑚 −25)
𝑄= 𝑃 𝑀 𝑆 𝑀 𝑋 𝐺 𝑆 𝑀 𝑋 𝐺 ∆𝐻 𝑟1 ° + 𝑆 𝑀 𝑋 𝐺 ∆𝐻 𝑟2 ° + 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝 𝑀 + 𝑆 𝑀 𝑋 𝐺 𝑐 𝑝 𝐷 + 𝑋 𝐺 𝑐 𝑝 𝑊 − 1− 𝑋 𝐺 𝑐 𝑝 𝐺 − 1− 𝑋 𝐺 𝑐 𝑝 𝐹𝐴 𝑇 𝑚 −25

32. Determination of Reactors Volumes Cost
Operation conditions:
Reactor Temperature = 255°C
Pressure, PT = bar
R = kJ.K/kmole

33. For CSTR;

35. The annual reactor cost;

37. LEVEL 4 SYNTHESIS OF CHEMICAL SEPARATION SYSTEM

38. Distillation Column

39. Sizing Distillation Column
Determination of Minimum Number of Stages
Minimum and Actual Reflux Ratio
𝑁 𝑚𝑖𝑛 = 𝑙𝑜𝑔 𝑑 𝐿𝐾 𝑑 𝐻𝐾 𝑏 𝐻𝐾 𝑏 𝐿𝐾 𝑙𝑜𝑔 𝛼 𝑚
𝑅 𝑚𝑖𝑛 = 𝑙𝑜𝑔 𝑥 𝐿𝐻𝑑 𝑑 𝑋𝐻𝐾𝑑 − 𝛼 𝐿𝐾,𝐻𝐾 ( 𝑋 𝐻𝐾𝑑 𝑋 𝐿𝐾 ) 𝛼 𝐿𝐾,𝐻𝐾 −1

40. Theoretical and Actual Number of Stages
The theoretical number of stages, N is calculated by using Gilliland correlation:
Calculated column diameter D = m
Column Height = m
𝑵− 𝑵 𝑴𝑰𝑵 𝑵+𝟏 =𝟏−𝒆𝒙𝒑 𝟏+𝟓𝟒.𝟒𝒙 𝟏+𝟏𝟏𝟕.𝟐𝒙 𝒙−𝟏 𝒙

41. Calculation for Distillation Column
Component
Feed
Distillate
Bottom
Molar flow
Mol fraction
Distearin
0.1648
0.0166
0.0009
0.2092
Glycerin
0.2408
0.9973
2.3643
0.0359
Monostearin
0.4309
0.0051
0.0003
0.5476
Fatty Acid
0.1635
0.0260
0.0015
0.2073
Total
1.0000

42. Fenske ( Nmin) Parameter/Component Glycerol (LK) Monostearin (HK)
Distillate Flow Rate, di
17.81
0.01
Bottom Flow Rate,bi
2.36
36.10
(αlk,hk)N
(αlk,hk)1
Nmin 14

43. Gilliland correlation Calculated column diameter D = 4
Gilliland correlation Calculated column diameter D = m Column Height = m
Rmin
1.75
Reflux Ratio, R
2.1 X
0.1129 Y N 28

44. Cost of Distillation Column
Where;
A = capacity or size parameter of the equipment
K1, K2, K3 = values used in the correlation
𝑙𝑜𝑔 10 𝐶 𝑝 𝑜 = 𝐾 1 + 𝐾 2 𝑙𝑜𝑔 10 𝐴 + 𝐾 3 [ 𝑙𝑜𝑔 10 𝐴 ] 2

45. EP4 = EP3 - 𝐶 𝑝 𝑜 (distillation column)

47. LEVEL 5 HEAT INTEGRATION

48. Heat Exchanger Network
Stream
Type
Tsupply (K)
Ttarget (K)
Total Heat Capacity Flowrate, FCp (KW/K)
Enthalpy Change, ∆H (KW) H1
Hot
498.15
373.15
8.76 H2
328.15
2.37 C1
Cold
298.15
0.834
25.011 C2
393.15
5.494
Total Q available = KW
Total Q that must be absorbed = KW

49. Temperature Intervals
Shifted temperature for the hot and cold stream in Pinch Technology
Stream
Type
Tsupply(K)
Ttarget(K)
TsS
TsT ∆T ∆H
FCp (KW/K) H1
Hot
498.15
373.15
493.15
368.15
-125
8.756 H2
328.15
323.15
-170
2.370 C1
Cold
298.15 30
25.011
0.834 C2
393.15 65
5.494

50. Temperature Intervals

51. Heat transfer to and from utilities for each temperature interval

52. CHAPTER 3 SIMULATION

54. CHAPTER 4 MATERIAL AND ENERGY BALANCE

55. Streams
Manual Calculation
(kg/hr)
Simulation
Error Percentage
(%) 1
0.00 2 3
0.22 4 5 6 7 8 9
4.71 10 11
15.834
12.79 12
4.62

56. Streams Manual Calculation (kg/hr) Simulation Error Percentage (%)13
0.58 14
1.29 15 16
0.08 17
0.44 18
0.65 19 20 21 22
0.06 23
14.75 24
1.48

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