Episode 47 : CONCEPTUAL DESIGN OF CHEMICAL PROCESSES

SAJJAD KHUDHUR ABBASCeo , Founder & Head of SHacademyChemical Engineering , Al-Muthanna University, IraqOil & Gas Safety and Health Professional – OSHACADEMYTrainer of Trainers (TOT) - Canadian Center of Human Development

Episode 47 : CONCEPTUAL DESIGN OF CHEMICAL

PROCESSES

INTRODUCTION Chemical process design is the application of chemical

engineering knowledge (chemical, physical and/or biological transformations of raw materials) into products and economics in the conceiving a chemical process plant to profitably manufacture chemicals in a reliable and safe manner without unduly affecting adversely the environment and society

Chemical process plants are by nature large capital investment projects that

are expensive to build and operate have very long life times and manufacture specific chemicals

Chemical process plants must be designed well to avoid large financial losses over long periods of times due to inefficient processes/poor operations

INTRODUCTION Main design objectives of chemical processes:

design of a grassroot plant ora retrofit design for existing chemical plants

Complimentary objectivesprofitable, safe, reliable, flexible, controllable

and operable Not all of these objectives can be fulfilled

however and some trade offs must be made in order to produce a practical design

UNIT OPERATIONS In the past, chemical process plants are

designed using unit operations first proposed by G.E. Davis in 1887

Unit operations was formalised by A.D. Little in 1915 as the defining principle of chemical engineering

The concept was earlier proposed by the ancient alchemists, in the course of transforming and purifying their chemicals through a series of operations of heating, distillation, evaporation etc.

UNIT OPERATIONSNew chemical process plants were then

designed byarranging the unit operations in the same

sequence as the original laboratory methods

increasing the size of equipment linearly for greater capacity

In the 40’s, it was realised that scaling-up is not linear and pilot plant studies needed to be done in order to determine the correct scaling-up parameters

UNIT OPERATIONS Up to the late 70’s, chemical process design

was still done byarranging unit operations in the sequence

proposed by the industrial chemists using block diagrams and later PFDs

performing the mass and energy balancesizing the individual equipmentdetermining the economic viability of the plant

Alternative PFDs were not easily generated due to

the empirical nature of the chemical technology the large number of uncertain variables to be

determined all at once

UNIT OPERATIONS Design parameters were determined in ad

hoc manner & specific for particular process

No systematic method for generating alternative PFDs and optimising them

Short cut methods of designing heat and mass transfer equipment already available

Equipment costing methods have been fairly developed using costing charts

Possible integration and optimisation of unit operations due to interconnections within the chemical process system was not understood

PROCESS SIMULATION With powerful computers and better

understanding of thermodynamics in the late 60’s to early 80’s, computational and optimisation methods were used in process system engineering

Since the 60’s, primitive process simulation softwares were owned by large petrochemical companies

These were mainly the sequential modular type where the unit operation modules were solved one by one in the direction of mass flow

PROCESS SIMULATION Modular simulation consists of

a top level of flowsheet topology where unit module are sequenced, recycle and tear streams determined, and convergence made,

a middle level where the unit operations are modeled and solved and

a lower level where physical and thermodynamic models are solved

By the late 70’s, the solution of modular flowsheets was significantly improved leading to simultaneous modular flowsheets which are the basis of commercial process simulation softwares such as

ASPEN/PLUS from Aspen Technology Inc. and HYSYS from Hyprotech Ltd

Most process simulations use phase equilibrium thermodynamic models including non-idealities in both liquid and gas phases for their unit operation models

Popular models activity

models used are the equation ofstatefor hydrocarbon mixtures and liquidcoefficients models for non-electrolyte,non-ideal solutions

Group contribution models such as UNIFAC are becoming popular when no empirical vapour-liquid equilibrium data is available

Rate-based models are very well developed and may well become more important when tray efficiency could not account for non-ideal behaviour

PROCESS SIMULATION

PROCESS SIMULATION In the 90’s, stoichiometric and

equilibriumreactor handling

models are primitive withpoor of multiple

reactions incompletely mixed and plug flow reactors Incorporation of rigorous generic models

for multi-phase industrial reactors is still a long way off

Some process simulator companies do model these reactors for individual process licence owners

PROCESS SIMULATION The generic modeling of adsorption, membrane

and solid drying processes are not well developed enough to be included in process simulations

A shortcut method for the generic design of adsorption columns presented by Wan Ramli Wan Daud 2000b shows some promise

Solids handling was neglected in process simulation work

It is now more important due to the increased popularity of fluidised bed reactors and pneumatic conveying

PROCESS SIMULATION In the 80’s and 90’s significant improvement

was made in the equation-oriented process simulation where the equations for all unit operations are combined and solved simultaneously

Allows specifications of certain design parameters without having to solve another iterative loop

The computational effort is reduced by the exploitation of sparse matrices

Succesful solution requires careful initialisation based on users’ past experience

It is used in quick on-line real time modelling and optimisation where models are simpler and initial points are taken from previous solutions

PROCESS SIMULATION Both simulations require simultaneous solution of

large sets of non-linear equations which are mainly based on Newton or quasi-Newton or Broyden methods due to their good convergence properties

Rapid solution of very large flowsheets can be achieved by a suitable decomposition strategy

by recycle tearing streams for the modular simulation by utilising powerful sparse matrix solvers for

equation oriented simulation Although process simulation is a powerful tool,

it is not possible to produce optimised design by simply using it because the optimum configuration and operating principle of the process plant could only be produced by process synthesis

PROCESS SYNTHESIS Contemporary process design method is

an iterative problem solving and optimisation method using both heuristic and algorithmic methods

Design method begins with the determination of the design requirements and objectives which are promulgated in either an economic or utlitarian way

A conceptual design is then produced through the synthesis of several feasible alternative designs and the rapid selection of the most viable of these alternatives based on an economic performance criterion without using rigorous performance models of their operational principles

PROCESS SYNTHESIS

PROCESS SYNTHESIS During synthesis, design variables or

parameters are selected or determined and optimised through

Heuristics, intuitions and experience or algorithmic methods using shortcut performance models

of the equipment or chemical process Complex design problems are decomposed

into their constituent parts where each part is further synthesised, its performance is modelled on its operational

principle and its design variables or parameters are

determined in a similar manner while maintaining integral relationship with

other parts as well as with the overall design

PROCESS SYNTHESIS First approach: Process synthesis can be solved

by mathematical modelling alone based on the principles of process flowsheeting

Assumes linearly

that technology emergesfrom which

science scientificis not true because

the physical phenomena in anknowledge onengineering artefact does not lead to knowledge on the operating principles and design of the artefact

The chemical process plant has a large number of variables that are defined by a smaller number of equations, with some inexplicable to deterministic models and most highly non-linear

Able to synthesise small plants where variables are defined adequately by equal number of equations unless efficient decomposition procedures are used

PROCESS SYNTHESIS

PROCESS SYNTHESIS Second approach: Process synthesis could

be solved by expert knowledge obtained from experience, intuition/insight and inspirations

Expressed as heuristic rules/rules of thumbs which set unknown parameters rapidly

Some heuristics relate external performance parameters with the operating variables of the artefact simply and directly without complex non-linear mathematical modelling

PROCESS SYNTHESIS The

synthesisproblem isdecisions forgenerating

decomposed into anand the

heirarchy ofexploring process alternatives starting fromtop down and considering a few designvariables at a time like peeling an onion

Basic assumption : design parameters at the top level also reflect design parameters further down

Old alchemical maxim of the relationship between the macrocosmos and the microcosmos: “what is above so below”

HEIRACHICAL PROCESS SYNTHESISFirst & Second Levels

In Douglas version, after decomposition by removing all the heat exchangers, the first level involves use of heuristics to select

Process Mode: Design Variables: Batch or continuous

The second level involves construction of the input-output structure of the process and targeting the production rate by using heuristics:

Whether the feed should be pretreated Destination of products Design variables : conversion of limiting

reactant and allowable purge concentration of excess reactant

Economic potential of process: Products sales less raw materials’ cost

HEIRACHICAL PROCESS SYNTHESISFirst & Second Levels

HEIRACHICAL PROCESS SYNTHESISInput-Output Structure

HEIRACHICAL PROCESS SYNTHESISInput-Output Structure: Destination StreamsToluene Hydro-Dealkylation Process

H , CH2 4

Toluene

Purge H , CH2

4Toluene Hydro-

Dealkylation Process

Benzene

Diphenyl

Component Normal Boiling Point(C)

Light/Heavy Destination

Hydrogen -253 Light Recycle dan PurgeMethane -161 Light Recycle dan PurgeBenzene 80 Heavy Main ProductToluene 111 Heavy RecycleDiphenyl 253 Heavy Fuel

HEIRACHICAL PROCESS SYNTHESISInput-Output Structure: Material BalanceToluene Hydro-Dealkylation Process

FG

FT

1 S

n B P

B2SB2

PB 1 SB

1 SB

PB

2SB

PB FE

SB 2SB

FH yFH FG FE

PM 1 yPH PG 1 yFH FG PB

S B F P SBBFH GPG FE 1 y

PH E Gy F P1 SB

PB

2SB

PG FG

1 1 y 1 S 2

PH By y

S

PBPH BFH

GF

PB n1 2n2

FT n1

FH FE n1 n2

PD n2

FT PB SB

PM FM

n1 n1 PB

SB

RG PG

Toluene Hydro-


PB

PD

HEIRACHICAL PROCESS SYNTHESISInput-Output Structure: Material BalanceToluene Hydro-Dealkylation Process

FG

FT

PB = 265 kgmole h-1 benzene

RG PG

Toluene Hydro-


PB

PD

XB FG

(kgmole h-1)FT

(kgmole h-1)PH

(kgmole h-1)PM

(kgmole h-1)PD

(kgmole h-1)0.1 312.49 266.13 31.31 281.75 0.560.2 312.64 266.35 31.33 281.99 0.680.3 312.84 266.67 31.37 282.31 0.830.4 313.13 267.12 31.42 282.77 1.060.5 313.58 267.81 31.50 283.49 1.410.6 314.34 268.99 31.63 284.70 1.990.7 315.82 271.27 31.90 287.06 3.130.8 319.51 276.97 32.55 292.94 5.980.9 336.48 303.20 35.56 320.02 19.10

HEIRACHICAL PROCESS SYNTHESISSecond Level Economic Potential

-5000000

Conversion of Limiting Reactant

5000000

0

10000000

15000000

20000000

25000000

30000000

0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9

Penukaran Toluena

hn)t/MR (i

onomk Einseot P

yph=0.1

yph=0.5

yph=0.65

yph=0.786

Econ

omic

Pot

entia

l (R

M)/Y

ear

Excess Reactant Concentration inPurge Stream

Toluene Hydro-Dealkylation Process

CB PB CFD PD CFP PG CT FT CH

FG

f PE

2

HEIRACHICAL PROCESS SYNTHESISInput-Output Structure: Destination Streams

Benzene Alkylation Process

Benzene Recycle

Propane and Propylene As Fuel

Cumene

P-diisopropyl Benzene As Fuel

Propylene

Benzene


Component Normal Boiling Point (C)


C3H8 -42.1 Light FuelC3H6 -47.8 Light FuelC6H6 80.1 Heavy RecycleC9H12 152.4 Heavy Main ProductC12H18

210.3 Heavy Fuel

HEIRACHICAL PROCESS SYNTHESISInput-Output Structure: Material Balance


PCFG

FB

PG

PD

RB

PC n1 n2

P 1 2 PP GF n n yP FB

n1PD n2

FP yFP FG PC

SC yPP PG

1 S

Pn C C

2SC

1 S P C C

1

2SC

n21 SP C

C CB 2S

F

1 S P C C

CD 2S

P

F Pr G P Pr Gy F yP

y 1 y 1 yPP

SC y FPPPFP

PC

GF


HEIRACHICAL PROCESS SYNTHESISInput-Output Structure: Material Balance


PCFG

FB

PG

PD

RB


P = 104 kgmole h-1

cumeneC

XB FG

(kgmole h-1)PD

(kgmole h-1)FB

(kgmole h-1)PG

(kgmole h-1)

0.1 113.04 0.00 103.99 9.040.2 113.10 0.03 104.02 9.050.3 113.17 0.06 104.05 9.050.4 113.24 0.10 104.09 9.060.5 113.33 0.13 104.13 9.070.6 113.41 0.17 104.17 9.070.7 113.51 0.22 104.21 9.080.8 113.61 0.27 104.26 9.090.9 113.73 0.32 104.31 9.101.0 135.75 10.45 114.44 10.86

HEIRACHICAL PROCESS SYNTHESISSecond Level Economic Potential

-5000000

45000000

40000000

35000000

30000000

25000000

20000000

15000000

10000000

5000000

00 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1


ypp=0.1 ypp=0.5 ypp=0.7 ypp=0.898

Excess Reactant Concentration inPurge Stream

Econ

omic

Pot

entia

l (R

M)/Y

ear

f PE 2 CC PC CFDIPB PDIPB CFP PP CP FP CB FB


HEIRACHICAL PROCESS SYNTHESISThird Level

The third level involves the construction of the reactor and recycle structures of the process by using heuristics to decide on

the number of reactor systems required their types (completely mixed or plug flow) operating modes and conditions and heat management number of recycle streams whether a gas recycle is required, and recycle flow rates as functions of conversion and mole

or recycle ratio Annual costs of reactors & compressors are

subtracted from economic potential at this level

HEIRACHICAL PROCESS SYNTHESISThird Level

HEIRACHICAL PROCESS SYNTHESISRecycle Structure


ReactorSeparation

& Purification System

Compressor

2 4

Benzene Product

Diphenyl Product

Hydrogen Feed

Toluene Feed

Toluene Recycle

Vapour Recycle RGH , CH

PurgeH , CH2 4


FT

T F R M

y F y

RFH G PHG X

yFH 1 1 yPH 1 SB 2 y y

S y

X

MR

R FH

PHTB PH

PB

G

RT 1 XT

FT

RT

Component Normal Boiling Point (C) Light/Heavy Destination

Hydrogen -253 Light Recycle dan PurgeMethane -161 Light Recycle dan PurgeBenzene 80 Heavy Main ProductToluene 111 Heavy RecycleDiphenyl 253 Heavy Fuel


Benzene Alkylation ProcessPropane & PropAs Fuel

ne

Propylene

Benzene

Benzene RB 1 XP MR yPF

FG

Reactor

Recycle RB

ylene

Separation& Purification

System

Cumene

P-diisopropyl Benze As Fuel

Component Normal Boiling Point (C)


C3H8 -42.1 Light FuelC3H6 -47.8 Light FuelC6H6 80.1 Heavy RecycleC9H12 152.4 Heavy Main ProductC12H18

210.3 Heavy Fuel

HEIRACHICAL PROCESS SYNTHESISAdiabatic Temperature

• For simple reaction A B,

• The adiabatic coversion

• Energy Balance for ReactorsN

j1• Adiabatic temperature

• In general

n j H Pc

T T F c

T Tm 0

k pk

m k 1

i1

KM

rj i pi am

m

n H Pc

T T 0 T T

n H

rj i pi

a i1

j1j

M

m a m jrj

Nm

Pc

i1

j1

M

i pi

Nm

Picpi FA 1 X A cPA FA

X AcpB i1

M FA 1 X A cPA FA X

AcpB

Tm 25

j1

FAcpA

F X H

N

n

H

oA Ar

mjr

X H

c c T

25X c c c

A pB pApA

pB pAm

orATa Tm

cpA Ta

Tm H

c c T

25X

pB pAa

or

Aa

Xa F c

T 25 nj

Hj1

i1

Na rj

M

i pia

800

600

1000

1200

1400

0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1Penukaran Toluena

kitabaid Auhu S

1800

1600Molar Ratio

MR=1

MR=2

MR=3

MR=4

HEIRACHICAL PROCESS SYNTHESISAdiabatic Temperature


Adi

abat

ic

Tem

pera

ture


HEIRACHICAL PROCESS SYNTHESISReactor Heat Management

XA

r = 1

r = 10r = 100

1.0

0T

Isothermal Reactor for Non-Autocatalytic Irreversible Reaction


Adiabatic Reactor for Non-Autocatalytic Irreversible Reaction

r = 1

r = 10 r = 100

XA

1.0

T

Endothermic Reaction

0

r = 1

r = 10 r = 100

XA

1.0

TExothermic Reaction

0


Endothermic Reaction Exothermic Reaction

Isothermal Reactor for Single Reversible Reaction

r = 1 r = 10 r = 100

XA

1.0

T0

r = 0

Equilbrium

r = 1

r = 10 r = 100

XA

1.0

T0

r = 0

Equilibrium


Adiabatic Reactor for Single Reversible Reaction

Endothermic Reaction Exothermic Reaction

Adiabatic Curve

r = 1 r = 10 r = 100

XA

1.0

T0

r = 0

Equilibrium

XAf

r = 1

r = 10 r = 100

XA

1.0

T0

r = 0

Equilibrium Adiabatic Curve

XAf

HEIRACHICAL PROCESS SYNTHESISCompressor and Reactor SizingToluene Hydro-Dealkylation Process

n 1

P P

n1 n

12 1

nZRT1

W RGT T P

P

n1 n2 1 2 1

FAo C Ao

0

rA

X A dX A V

A

V X A

F CAo Ao r

F ln1 1 X T

y F y R F oP H Go F H Gk exp(E / RT )V

0.5

XT Compressor (kW)

Reactor Volume (m3)

Reactor Length

(m)

Diameter (m)

0.1 3702.73 307.68 22.87 4.140.2 1787.160 314.99 22.87 4.190.3 1149.083 324.39 22.87 4.250.4 830.590 336.69 22.87 4.330.5 640.259 353.34 22.87 4.440.6 514.624 376.95 22.87 4.580.7 427.371 413.26 25.61 4.530.8 368.636 478.94 25.61 4.88

0.9 358.480 668.20 25.61 5.76

HEIRACHICAL PROCESS SYNTHESISThird Level Economic Potential

Toluene Hydro-Dealkylation Process2.11 F

Cp

81508.743IMSd

0.82

d

IMSkpt

W W K

Fm Fp FIR

MSd MSk

rt D L I

I 7775.3

K 1.066 0.82 2.183

Material of Construction

Carbon Steel Carbon steel chromium- molybdenum

Stainless Steel

Fm 1.00 2.15 3.75

Compressor Fd

Centrifugal compressor with electric motor 1.0Centrifugal compressor with turbine 1.15Reciprocating compressor with steam 1.07Reciprocating compressor with electric motor 1.29Reciprocating compressor with engine 1.82

Pressure (Bar)

1.6 6.8 13.6 20.4 27.2 34.0 40.8 47.6 54.4 61.2 68.0

FP 1.00 1.05 1.15 1.20 1.35 1.45 1.6 1.8 1.9 2.3 2.5


-20000000

-30000000

-40000000

-10000000

0

30000000

20000000

Molar Ratio10000000

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8

Conversion of Limiting Reactant0.9 1

Penukaran

3asr Aimo nko Esiento P

MR=2 MR=3 MR=4 MR=5

Econ

omic

Pot

entia

l (R

M)/Y

ear


CH FG K pt K rtf PE 3 C B PB C FD PD C FP PG CT FT



0

0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.9 1

-10000000Conversion of Limiting Reactant

f PE 2 CC PC C FDIPB PDIPB C FP PP C P FP C B FB K rt

10000000

20000000

40000000

Molar Ratio30000000

50000000

ypp=0.1

ypp=0.5

ypp=0.7

ypp=0.898

Econ

omic

Pot

entia

l (R

M)/Y

ear

HEIRACHICAL PROCESS SYNTHESISFourth Level

The fourth level involves the synthesis ofthe separation structure of the flow sheet

Reactor products are to be separated into a liquid and a vapor phase by cooling, decompressing

orboth with the main product in the liquid phasebecause liquid purification technologylike distillation can produce very pure product

The liquid stream is sent to the liquid separation train consisting usually of a distillation train that are sequenced using heuristics

The vapor stream may be purged and the rest of the vapor recovered &/or recycled, or be condensed into liquid, which is sent to the liquid separation train, and the rest ofuncondensible vapor recovered, recycled or purged


Selection of simple or complex columns and order of distillation columns sequence using heuristics

Distillation columns design using short cutmethodse.g. Fenske-Underwood-Gilliland (FUG)

For non-ideal and azeotropic distillation Identify azeotropes & alternative separators Select entrainers & identify feasible distillate &

bottom products compositions Design variables: pressure, temperature & product

recoveries at flash drum, absorbers, adsorbers, membrane modules & distillation columns

Annual costs of separation systems are added to the economic potential at this level


HEIRACHICAL PROCESS SYNTHESISSeparation Structure

HEIRACHICAL PROCESS SYNTHESISSeparation Structure: Sub-cooled Liquid

HEIRACHICAL PROCESS SYNTHESISSeparation Structure:Both Superheated Vapor & Sub-cooled Liquid

HEIRACHICAL PROCESS SYNTHESISSeparation Structure: Superheated Vapor

HEIRACHICAL PROCESS SYNTHESISFlashing to Separate Liquid and Vapour

Dew point

Bubble point

Flash calculation using Rachford-Rice method

Ki 1 xi

yi

1 K1

ii

(3.23) yi Ki xi 1

y

K i zi

1 K 1

zi K i

xi

zi

ii

x

1 ix

(3.24)

(3.25)

f

1 Ki zii 1 1 Ki

1

C

i 1

C

i 1

0

C

i iy x


Top Product

Toluene Hydro-Dealkylation Process3.3 bar dan 35CReactor Product

Toluene Conversion

XT

Toluene (kgmole h-1)

Methane (kgmole h-1)

Hydrogen (kgmole h-1)

0.1 2395.15 19562.99 13042.750.2 1065.41 9591.88 6395.4890.3 622.22 6270.52 4181.460.4 400.67 4612.69 3076.540.5 267.81 3622.01 2416.550.6 179.32 2968.12 1981.400.7 116.26 2514.08 1680.230.8 69.24 2208.71 1480.450.9 33.69 2157.52 1463.81

Toluene Conversion

Methane (kgmole h-1)

Hydrogen(kgmole h-1)

Benzene(kgmole h-1)

Toluene(kgmole h-1)

Diphenyl(kgmole h-1)

XT

0.119489.20 12986.6 17.24 51.29 1.8x10-5

0.2 9555.07 6367.45 16.96 22.43 2.1x10-5

0.3 6246.13 4162.88 16.75 12.93 2.6x10-5

0.4 4594.45 3062.64 16.48 8.19 3.2x10-5

0.5 3607.44 2405.44 16.22 5.38 4.2x10-5

0.6 2955.99 1972.15 15.98 3.55 5.9x10-5

0.7 2503.70 1672.30 15.82 2.28 9.2x10-5

0.8 2199.56 1473.44 15.77 1.35 1.7x10-4

0.9 2149.03 1457.23 16.56 0.69 5.9x10-4

Bottom ProductToluene

Conversion XT

Metahne(kgmole h-1)

Hydrogen(kgmole h-1)

Benzene(kgmole h-1)

Toluene(kgmole h-1)


0.1 73.79 56.20 247.76 2343.86 0.56360.2 36.81 28.04 248.04 1042.99 0.67660.3 24.39 18.58 248.25 609.29 0.83250.4 18.25 13.90 248.52 392.49 1.05800.5 14.58 11.11 248.78 262.43 1.40570.6 12.13 9.25 249.025 175.78 1.99250.7 10.39 7.93 249.18 113.98 3.13310.8 9.16 7.01 249.235 67.89 5.98270.9 8.49 6.58 248.44 33.00 19.0978



Top Product

1.75 bar dan 90C (a)

1.75 bar dan 90CReactor Product

Propylene Conversion

XP

Propylene (kgmole h-1)

Propane (kgmole h-1)

Benzene (kgmole h-1)

DIPB(kgmole h-1)

0.1 94.121 5.273 1975.874 00.2 83.711 5.276 936.505 0.0300.3 73.292 5.279 590.055 0.0620.4 62.863 5.283 416.837 0.0960.5 52.423 5.286 312.912 0.1330.6 41.971 5.290 243.635 0.1740.7 31.505 5.295 194.157 0.2180.8 21.023 5.300 157.055 0.2660.9 10.522 5.305 128.205 0.320

0.99 1.256 6.333 137.868 10.451


XP




Cumene(kgmole h-1)

0.1 72.4 3.905 195.612 1.280.2 72.1 4.441 159.206 2.350.3 64.7 4.575 118.011 2.860.4 55.7 4.597 85.452 2.950.5 46.1 4.562 61.018 2.770.6 36.3 4.478 42.636 2.440.7 26.4 4.316 28.1527 1.960.8 16.4 3.991 16.491 1.360.9 6.76 3.218 7.180 0.693

0.99 7.19 0.313445 0.2760.0236Bottom Product


XP




Cumene(kgmole h-1)

DIPB(kgmole h-1)

0.1 21.748 1.368 1780.262 102.717 00.2 11.622 0.835 777.299 101.639 0.0290.3 8.551 0.704 472.044 101.134 0.0610.4 7.136 0.686 331.385 101.047 0.0960.5 6.289 0.725 251.894 101.221 0.1330.6 5.653 0.812 200.999 101.557 0.1730.7 5.134 0.979 166.004 102.037 0.2170.8 4.617 1.309 140.564 102.632 0.2660.9 3.762 2.087 121.025 103.300 0.319

0.99 1.184 6.019 137.592 103.970 10.450

HEIRACHICAL PROCESS SYNTHESISSeparation Structure: Liquid Separation

HEIRACHICAL PROCESS SYNTHESISSequencing of Simple Distillation Columns

Direct Sequence Lightest First

Indirect Sequence Heaviest First

1, 2, 3

2, 3

1 2

3

1, 2, 3

1, 2

2

1

3

HEIRACHICAL PROCESS SYNTHESISSequencing of Complex Columns

Complex Columns: Common Reboiler

Complex Columns Common Condenser

1, 2, 3

1 2

3

1, 2, 3

1

32

HEIRACHICAL PROCESS SYNTHESISSequencing of Complex Columns

Complex Columns:Both Top & Bottom Products of 1st Column as Feeds to 2nd Column withOne Side Product

Complex Columns: Side Product Above or Below Feed Point

1, 2, 3

3

1

2 1, 2, 3

1

3

2

1, 2, 3

1

2

3

HEIRACHICAL PROCESS SYNTHESISShort-Cut Method for Multi-component Distillation

Fenske-Underwood-Gilliland (FUG)• Fenske Equation to estimate minimum number of theoretical plate

• Underwood Equation to estimate minimum reflux ratio

• Gilliland Equation to Estimate number of theoretical plates

• Plate Efficiency: O’Connel Correlation

• Area of Condenser

• Area of Reboiler

1 LK HK

ln LK 1 HK

min ln

m

N

LK , HK LK ,

HK N

1 2

1 m

xD,LK xF ,LK LK / HK xD,HK xF ,HK

1LK / HK

Rmin R 1.2Rmin

0.5688 N Nmin min 0.75

1N 1 R 1

R R

N 2 N m

Eo

0.25

2

2.841

F

A U T

T

T

T T

ln dewc cwi

cwo dewc TVH v

cwic cwoc

U T T

R s dewR

VH vRA

HEIRACHICAL PROCESS SYNTHESISShort-Cut Method for Multi-component Distillation

• Height of distillation tower

• Diameter by using Fair Correlation for

H 0.69N Eo

0.01

0.1

1

0.01 0.1 1 10

F LV

Cf

(m/s

)

0.127 m

0.229 m0.305 m

0.610 m0.457 m

0.914 m

Distance between plate

0.5

L V V L

LM F

LV VM

C FST FF FHACF

FFT = (L/20)0.2

FF < 0.75

0.6u

1 AA

1 2

4VM V

V

D df

T

FHA = 1 if Ah/Aa > 0.1

FHA = 5(Ah/Aa) + 0.5if 0.06 > A /A > 0.1h a

1 2

L

V

V

u

C

f

HEIRACHICAL PROCESS SYNTHESISPreliminary PFD without Heat ExchangersToluene Hydro-Dealkylation Process

HEIRACHICAL PROCESS SYNTHESISPreliminary PFD without Heat ExchangersToluene Hydro-Dealkylation ProcessDesign of stabilizer Top Product

XTRmin R Hydrogen

(kgmole h1)Methane

(kgmole h-1)Benzene

(kgmole h-1)Toluene

(kgmole h-1)

0.1 0.354 0.530 56.194 73.780 2.478 0.0001230.2 0.346 0.520 28.036 36.807 2.480 6.1x10-5

0.3 0.341 0.511 18.577 24.386 2.482 4.1x10-5

0.4 0.335 0.502 13.902 18.246 2.485 3.0x10-5

0.5 0.328 0.493 11.109 14.576 2.488 2.4x10-5

0.6 0.323 0.484 9.253 12.133 2.490 2.0x10-5

0.7 0.318 0.477 7.930 10.387 2.492 1.7x10-5

0.8 0.313 0.469 7.012 9.158 2.492 1.5x10-5

0.9 0.309 0.464 6.581 8.491 2.484 1.4x10-5

Bottom ProductXT Hydrogen

(kgmol j-1)Benzene(kgmol j-1)

Toluene(kgmol j-1)

Diphenyl(kgmol j-1)

0.1 0.0056 245.2841 2343.8599 0.56360.2 0.0028 245.5596 1042.9852 0.67660.3 0.0019 245.7695 609.2914 0.83250.4 0.0014 246.0318 392.4874 1.05800.5 0.0011 246.2942 262.4312 1.40570.6 0.00093 246.5303 175.7762 1.99250.7 0.00079 246.6877 113.9812 3.13310.8 0.00070 246.7402 67.8907 5.9827

0.9 0.00066 245.9531 32.9967 19.0978

XT Height (m)

Diameter (m)

Condenser Area (m2)

Reboiler Area (m2)

0.1 23.25 0.399 739.49 41.630.2 23.25 0.284 111.29 20.020.3 23.25 0.233 37.97 12.900.4 23.25 0.202 21.83 8.060.5 23.25 0.182 14.60 6.250.6 23.25 0.167 10.95 5.130.7 23.25 0.156 8.65 4.440.8 23.25 0.147 7.15 4.23

0.9 23.25 0.142 6.46 5.86

HEIRACHICAL PROCESS SYNTHESISPreliminary PFD without Heat ExchangersToluene Hydro-Dealkylation ProcessDesign of benzene tower Top Product

XT RminR Benzene

(kgmole h-1)Toluene

(kgmole h-1)0.1 7.898 11.846 245.210 0.701

0.2 3.845 5.767 245.486 0.312

0.3 2.503 3.754 245.696 0.182

0.4 1.835 2.752 245.958 0.117

0.5 1.436 2.154 246.220 0.078

0.6 1.173 1.759 246.456 0.053

0.7 0.990 1.486 246.614 0.034

0.8 0.864 1.296 246.666 0.020

0.9 0.799 1.199 245.879 0.0098

Bottom Product

XT Benzene(kgmole h-1)

Toluene(kgmole h-1)


Minimum no. of plates

No. of theoretical

plates

No. of actual plates

0.1 0.0736 2343.159 0.5636 19.1 38.2 59

0.2 0.0737 1042.673 0.6766 18.9 37.7 59

0.3 0.0737 609.109 0.8325 18.6 37.2 60

0.4 0.0738 392.370 1.0580 18.4 36.8 60

0.5 0.0744 262.353 1.4056 18.2 36.4 61

0.6 0.0740 175.724 1.9925 18.0 36.0 62

0.7 0.0740 113.947 3.1331 17.8 35.6 62

0.8 0.0740 67.870 5.9827 17.7 35.4 63

0.9 0.0738 32.987 19.0978 17.7 35.3 64

XT Height (m)

Diameter (m)

Condenser Area (m2)

Reboiler Area (m2)

0.1 44.2 3.3 1291.64 327.610.2 44.8 2.4 700.02 173.110.3 45.4 2.0 494.61 122.700.4 45.9 1.8 390.68 98.080.5 46.4 1.6 328.70 84.060.6 46.8 1.5 287.83 76.080.7 47.3 1.4 259.42 73.150.8 47.7 1.4 244.89 76.85

0.9 47.8 1.3 257.89 110.75

HEIRACHICAL PROCESS SYNTHESISPreliminary PFD without Heat ExchangersToluene Hydro-Dealkylation ProcessDesign of toluene tower Top Product

XT RminR Toluene

(kgmole h-1)Diphenyl(kgmole h-1)

0.1 0.0413 0.0621 2342.456 0.000168

0.2 0.0414 0.0621 1042.361 0.000202

0.3 0.0414 0.0621 608.927 0.000249

0.4 0.0415 0.0622 392.252 0.000316

0.5 0.0416 0.0624 262.274 0.000420

0.6 0.0418 0.0627 175.671 0.000596

0.7 0.0425 0.0637 113.913 0.000936

0.8 0.0450 0.0675 67.850 0.001788

0.9 0.0653 0.0979 32.977 0.005708

Bottom Product

XT Toluene(kgmole h-1)


Minimum no. of plates

No. of theoretical

plates

No. of actual plates

0.1 0.703 0.563 5.04 10.08 15

0.2 0.313 0.676 5.04 10.08 16

0.3 0.183 0.832 5.04 10.08 16

0.4 0.118 1.058 5.04 10.08 17

0.5 0.079 1.405 5.04 10.08 17

0.6 0.053 1.992 5.04 10.08 17

0.7 0.034 3.132 5.04 10.08 18

0.8 0.020 5.981 5.04 10.08 18

0.9 0.009 19.092 5.04 10.08 18

XT Height (m)

Diameter (m)

Condenser Area (m2)

Reboiler Area (m2)

0.1 14.6 5.7 645.60 404.20

0.2 14.9 3.8 287.29 231.79

0.3 15.2 2.9 167.84 169.53

0.4 15.5 2.3 108.12 137.31

0.5 15.7 1.9 72.31 114.46

0.6 16.0 1.6 48.45 95.13

0.7 16.2 1.3 31.45 73.52

0.8 16.5 1.0 18.80 49.22

0.9 16.5 0.7 9.40 25.89

4asr Ai

onomk Eiensot P

-20000000

-25000000

-30000000

-35000000

-40000000

-45000000


-15000000

-10000000

-5000000

0

20000000

15000000

10000000

5000000

0 0.2 0.4 0.6 0.8 1

Penukaran

yph=0.4

yph=0.1

yph=0.2

yph=0.3

HEIRACHICAL PROCESS SYNTHESISFourth Level Economic Potential

Econ

omic

Pot

entia

l (R

M)/Y

ear

Molar Ratio

HEIRACHICAL PROCESS SYNTHESISFifth Level

In the fifth level, need for heat exchanges is reconsidered Heat exchanger network (HEN) is optimized & integrated by pinch

analysis based on First & Second Law of Thermodynamics Targeting for minimum number of heat exchangers (Fisrt Law) and

minimum utility requirement (Second Law) Identification of Hot & Cold Streams Second Law: Minimum approach temperature difference: 10C First Law: Energy cascade diagram Second Law: Temperature-enthalpy & grand composite

curves: Identification of pinch temperature HEN synthesis above & below pinch temperature Optimization of HEN synthesis by stream splitting & removal

of loops

HEIRACHICAL PROCESS SYNTHESISFifth Level

HEIRACHICAL PROCESS SYNTHESISPreliminary PFD with Heat Exchangers

HEIRACHICAL PROCESS SYNTHESISPreliminary PFD with Heat Exchangers

HEIRACHICAL PROCESS SYNTHESISHot & Cold Streams; Energy Cascade Diagram

Temp. Int. 2 = 40C

Temp. Int. 1 = 50C

Temp. Int. 3 = 10C

Temp. Int. 4 = 30C

Temp. Int. 5 = 20C120 110

150 140

190200

250 240

100 90

FCp 1000 W/C 4000 W/C 3000 W/C 6000 W/C

1

2

3

4

160 150

C C

0 W

C250

C240

200 190

150 140

100 90

Cold Utility

W

70,000 W 10,000 W

-40,000 WTemp. Int. 2

-80,000 WTemp. Int. 3

20,000 WTemp. Int. 5

W

60,000 W

Hot Utility

Stream No.

Stream Condition

Stream Enthalpy / C

Tin

(C)Tout

(C)

1 Hot 1000 250 1202 Hot 4000 200 1003 Cold 3000 90 1504 Cold 6000 130 190

To tal


50,000


40,000

HEIRACHICAL PROCESS SYNTHESISTemperature-Enthalpy & Grand Composite Curves

90

110

130

150

170

190

210

230

250

0 100 200 300 400

Enthalpy (kW)

500 600

Entalpi (kW)

C)o(

uhu S

Entalpi PanasEntalpi SejukEntalpi Sejuk Teranjak

90

110

130

150

170

190

210

230

250

0 20 40 60

Enthalpy (kW)

80 100 120

Entalpi (kW)

C)o(

uhu S

Hot Enthalpy

Cold Enthalpy

Shifted Cold Enthalpy

Tem

pera

ture

(oC

)

Tem

pera

ture

(oC

)

HEIRACHICAL PROCESS SYNTHESISIntegrated PFD

H I

EIRACHICAL PROCESS SYNTHESISntegrated PFD

HEIRACHICAL PROCESS SYNTHESISSixth Level

Poor process static & dynamic properties arise from usingeconomic viability for process selection causing off-spec products & excessive utilities

Seider et al and Daud (2001) added a sixth level, where a plant-wide control scheme is developed by using heuristics first introduced by Newell and Lee

Selection of Control Variables: Heuristic 1: Select state variable representing inventory

that is not self regulating Heuristic 2 Select state variable representing self regulating

inventory that transgress equipment’s limit or process condition

Heuristik 3 Select state variable representing self regulating inventory that interacts with another inventory

Selection of Manipulated Variables: Heuristic 1: Select variable that acts directly with control variable Heuristic 2: Select variable that is more sensitive to control variable changes Heuristic 3: Select variable that acts vary fast Heuristic 4: Select variable that does not

interact with other control loops Heuristic 5: Select variable that does

not recycle any disturbance

HEIRACHICAL PROCESS SYNTHESISSixth Level

HEIRACHICAL PROCESS SYNTHESISMass & Energy Inventory Control: Reactor

L

LT

LCR

T

TT

FCR FT

TCR

HEIRACHICAL PROCESS SYNTHESISMass & Energy Inventory Control: Heater

HEIRACHICAL PROCESS SYNTHESISDistillation Control: Cut Control Top Product

LCR2

FCR1LCR1

PCR1

R1FI1

FT1

L

LT1 P

PT1

L

LT2

FT1

FT2

R2 FCR2

HEIRACHICAL PROCESS SYNTHESISDistillation Control: Cut Control Bottom

FCR2

PCR1

R1FI1

FT1

L

LT1 P

PT1

L

LT2

FT3

R2

FT2

FCR1

LCR1

LCR2

HEIRACHICAL PROCESS SYNTHESISDistillation Control: Product Quality Control

PCR1

L

LT1 P

PT1

L

LT2

F

QT2

TCR1

FT1

T

TT1

F

QT1

QCR1

LCR1

LCR2

FCR1

QCR2

PROCESS SYNTHESIS

PROCESS SYNTHESIS & OPTIMISATION

The third approach is the algorithmic method to search for and optimise process alternatives

Process synthesis involving heavy mathematical modelling are decomposed efficiently due to very large combinatorial flowsheet possibilities and then optimised

One approach is a tree search in the space of design decisions where design decisions are recorded at a node which can be backtracked to a previous node & branched in different directions

The solution is optimised by using mixed integer linear programming (MILP)


Another method isthe

creation ofasuperstructure of decisions containing most

if not all design alternatives and then using mixed integer non linear programming (MINLP) to optimise them

Large superstructures might lead to very large MINLP problems that might be unsolvable

A viable alternative is to reduce the process alternatives through the use of heuristics and then optimise the reduced superstructure using MINLP or MILP


The most popular non linear programming algorithm used in process optimisation is the successive quadratic programming (SQP)requires less function evaluationsdoes not require feasible

points at intemediate iterations andconverges to an optimal solution from an

infeasible point.


Optimisation of reactor networks is not very well developed mainly due to the non-linear characteristics of reacting systemsDifficult to infer heuristic rules andDifficult to converge

algorithmic methodsNovel method proposed by Glasser et

al. 1987 is to plot an attainable region consisting of all the family of reactor network solutions


It is sufficient to get the reactor network at the boundary of the attainable region because any interior point is simply the mixture of the boundary points

In two dimensional problems, the reactors need to be continuous stirred tank reactors (CSTR) and plug flow reactors (PFR) only

The remaining problem is the integration of reactor networks with the separation system

CURRENT AND FUTURE DEVELOPMENT

More efforts should be devoted to the generic modelling of adsorption membrane solid drying solids handling especially fluidisation

and pneumatic conveyingFurther work on integrating of

process control and process synthesis should be developed using the structural control matrix approach


Important issues being neglected are safe design and operation and waste minimisation

Heuristic approach of Kletzusing

keywordslike intensification, substitution, and attenuation pioneered chemical process plant design for safety

Recentlyrapid inhenrentlysafe 2000

risk analysis is used todesign by Khan &Abbasia

A related issue is design for maintainability


The minimum addition of chemical species and their minimum productionandrejection pioneered minimum

in the massexchangenetwork by El-Halwagi using thenumber of “mass exchangers”

can minimise wastes Flower et al first proposed the use of mass

exchange networks for waste minimisation Recently Noureldina & El-Halwagi

reported a mass exchange network-based method for pollution prevention


A method proposed recently by Dantus & Higha is to evaluate source reduction alternatives byeconomic performance including

waste related costs in an environmental accounting framework and

the environmental impact of the alternative


A new method which is now becoming the trend is the combination ofeconomic objectives and life cycle assessment (LCA)-based

environmental objectives Usesgoal programming to identify the Pareto

surface of non inferior solutions Moreresearch incorporating environmental

should bedirected at waste

minimisation and impact ideas

in theheuristics-based method of Douglas

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