Benzene Chlorination Case Study∗
Jeffrey D. Ward, Duncan A. Mellichamp, Michael F. Doherty
Department of Chemical Engineering, University of California
Santa Barbara, California 93106-5080
1
Introduction
This document presents a case study which illustrates the analysis developed by Ward, et al.
Chlorination reactions are employed to introduce reactive sites on organic molecules. For example, one route in the production of phenol (C6 H5 OH) from benzene is via the intermediate
chlorobenzene (C6 H5 Cl). In this case care must be taken to minimize the production of higher
chlorinated benzenes. The reactions are:
benzene + Cl2 → chlorobenzene + HCl
chlorobenzene + Cl2 → dichlorobenzene + HCl
These reactions can be carried out in a CSTR in the liquid phase at 60 ◦ C. Chlorine gas is introduced
into the reactor and dissolved in the liquid by means of a sparger. We assume that dichlorobenzene
has no value and must be disposed of safely, i.e. the second reaction is a representative byproduct
forming reaction. The reactor temperature and the per-pass conversion of benzene are kept low in
order to suppress the undesired reaction. Unreacted chlorine, hydrogen chloride, and catalyst are
removed from the reactor effluent stream through stripping operations which are not considered
in this case study. All three isomers of dichlorobenzene (ortho-, meta-, and para-) are expected
to be formed, however these species have very similar normal boiling points and for practical
purposes will come out together when separation is accomplished by distillation. Therefore we
treat dichlorobenzene as one species. The setup of this case study is similar to one developed by
Kokossis and Floudas.1
∗
This case study is a web-published supplement to the paper entitled “The Importance of Process Chemistry in
Selecting the Operating Policy for Plants With Recycle” by Ward, Mellichamp, and Doherty.
1
Silberstein, et al.2 report kinetic data for the above reactions catalyzed homogeneously by
stannic chloride. The reactions are assumed to be third order overall, being first order in the
concentration of catalyst, chlorine, and benzene (or chlorobenzene):
r0 = k0† [benzene] [Cl2 ] [SnCl4 ]
(1)
r1 = k1† [chlorobenzene] [Cl2 ] [SnCl4 ]
(2)
The concentration of chlorine in the liquid is determined primarily by mass transfer and solubility limitations and is assumed not to be available as a degree of freedom in the process design or
operation. Furthermore, because the catalyst concentration and chlorine concentration influence
both reaction rates in the same manner, their value does not influence the selectivity-conversion
profile. Therefore, it is assumed that these parameters are fixed and not available as degrees of freedom: [SnCl4 ] = 0.030 mole/L and [Cl2 ] = 0.25 mole/L. The reactor temperature is kept constant
at 60
◦
With these assumptions, the reaction rates are given by:
r0 = k0 [benzene]
k0 = 0.22hr−1
r1 = k1 [chlorobenzene] k1 = 0.041hr−1
With these simplifications, the process chemistry is effectively of the type A→B→C, which is nonbounded.
Table 1 shows properties of the species. Separation is accomplished by a series of distillation
columns. Candidates for the distillation sequence include the direct split (A/BC followed by B/C)
and the indirect split (AB/C) followed by (A/B). The direct split is selected following the heuristic
that for controllability it is desirable to minimize the number of distillation columns in the recycle
loop. Furthermore, because it is anticipated that the per-pass conversion of benzene will be low
(less than 0.5), unreacted benzene will be the primary component in the reactor effluent stream, and
the direct split requires that benzene be boiled up only once compared to twice with the indirect
split. Therefore, the direct split is expected to have a lower operating cost. The reactor network
considered in this case study is a single CSTR, although a PFR or a cascade of CSTRs would be
expected to give a better selectivity-conversion profile. Thus the process flowsheet at Douglas’3
level 4 is as shown in Figure 1.
2
Calculations
Following the method of Ward, et al., expressions can be developed for the flowrates of the process
streams in Figure 1 and the reactor volume as a function of the principle design degree of free2
dom, the recycle flow rate of benzene RB . Subscripts B, C, D, and Cl refer to species benzene,
chlorobenzene, dichlorobenzene, and chlorine, respectively.
The byproduct production rate is given by:
PD (RB ) =
1
k1 PC2
PC
k0 RB 1 − k10 R
B
(3)
and by global material balance the fresh feed streams are given by:
FF,B (RB ) = PC + PD (RB )
(4)
FF,Cl (RB ) = PC + 2PD (RB )
(5)
Furthermore, the reactor volume is given by:
V (RB ) = [PC + PD (RB )]
q(RB )
k0 RB
(6)
where:
q(RB ) = vB RB + vC PC + vD PD (RB )
3
(7)
Process Design
The nominal process is designed to produce 50 kmol/hr of chlorobenzene (47 million kg/yr) with a
reactor temperature of 60 ◦ C. Distillation columns were sized and costed following the methods of
Doherty and Malone.4 Figure 2 shows the economic potential of the process at level 4 as a function
of the key design variable, the recycle flow rate of benzene RB . It is important to note that what
is plotted in this figure is the actual shortcut estimate of the profitability of the process not the
simplified cost function C of Ward, et al. As expected, because benzene is a non-bounded species,
there is an optimum in the economic potential away from the constraints. The global maximum
economic potential of the nominal process is with the recycle flow rate of benzene equal to 220
kmol/hr, corresponding to an economic potential of $3.3 million/year. Table 2 shows the stream
table corresponding to this design.
In order to verify the calculations, the benzene chlorination process was simulated in HYSYS.5
Table 3 shows the stream table for the HYSYS simulation, and Table 4 compares several properties
of the process equipment determined by shortcut calculations and by HYSYS. The two stream
tables show excellent agreement. The properties of the unit operations also show good agreement.
The somewhat larger reactor volume suggested by the HYSYS simulation is due in large part to
3
the fact that the volume occupied by the chlorine in the reactor is neglected in the shortcut design.
The difference in heat duty in the first column between the shortcut calculation and the HYSYS
simulation is due to the use of an average value of the molar enthalpy of vaporization in the shortcut
calculations.
In order to make the process operable, it is necessary to overdesign the actual process equipment
by some amount compared to the nominal design. This is accomplished by sizing the process
equipment for a production rate 30% greater than the nominal value. The inoperable process with
smaller process equipment sizes is hereafter called the “nominal process,” while the flexible process
with larger equipment sizes is called the “flexible process.” Table 5 compares some of the equipment
sizes for the nominal process and the flexible process.
4
Operating Policies
Figure 3 shows the optimization problem for the flexible process after it has been built. The solid
line shows the operating economic potential of the flexible process as a function of the operational
degree of freedom, the recycle flow rate of benzene. The operating economic potential of the plant
after it has been built includes the fixed capital cost associated with the process as well as the
variable operating costs. The vertical dashed lines show the reactor volume constraint (on the
left) and the recycle capacity constraint (on the right). Again as expected, there is a maximum
in the economic potential function. However, the location of the maximum has shifted now that
capital costs are fixed, and lies outside the feasible region. As is sometimes the case for nonbounded chemistries, the optimal operating point lies on the recycle capacity constraint, at 290
kmol/hr. This operating point corresponds to a reactor holdup of 72% of the maximum value. The
economic potential of the process at this point is $3.1 million/year. Note that this is somewhat less
than the maximum achievable economic potential at the design stage, because this value reflects
the additional cost for oversizing the process equipment. An alternative, inferior operating policy
would be to operate with the reactor completely full. The economic potential of the process with
this inferior operating policy is $2.2 million/year, a loss of nearly 30% compared to the maximum
achievable economic potential.
Now consider a decrease in the production rate of 50%. The optimization problem is shown
in Figure 4. Because much less product is being produced, the economic potential of the process
is significantly reduced. Also, the optimal recycle flow rate has been reduced by 50%, and now
4
lies within the process constraints, at 202 kmol/hr, where the economic potential of the process
is $494 thousand/year and the reactor holdup is at 34% of the maximum value. The process is
within the region of operation where the optimal operating policy is to scale the reactor holdup
and recycle flow rate linearly with production rate. The process incurs a loss of $1.6 million/year
if it is operated with the reactor completely full. The economic potential of the process when it is
operated on the recycle capacity constraint is $458 thousand/year, corresponding to a loss of 7.3%
relative to the maximum achievable economic potential.
References
[1] Kokossis, A C.; Floudas, C. A. Synthesis of Isothermal Reactor-Separator-Recycle Systems.
Chem. Eng. Sci. 1991, 46, 1361.
[2] Silberstein, B.; Bliss, H.; Butt, J. B. Kinetics of Homogeneously Catalyzed Gas-Liquid Reactions: Chlorination of Benzene with Stannic Chloride Catalyst. Ind. Eng. Chem. Fund. 1969,
8, 366.
[3] Douglas, J. M.; Conceptual Design of Chemical Processes; McGraw-Hill: NewYork, 1988.
[4] Doherty, M. F.; Malone, M. F. Conceptual Design of Distillation Systems; McGraw Hill: New
York, 2001.
[5] HYSYS v. 3.1 (build 4815) Hyprotech Ltd. 2002.
[6] DISTIL v. 5.0 (build 4696) Hyprotech Ltd. 2001.
[7] CRC Handbook of Chemistry and Physics. CRC Press: Boca Raton, FL, 1995.
5
Table 1: Properties of Species
Benzene
Chlorine
Chlorobenzene
Dichlorobenzene
Value ($/kmol)
15.00
15.00
45.00
0.00
molecular weight (g/mol)
78.11
70.91
112.56
147.00
normal boiling point (◦ C)
80.1
−34.06
130
173-180
mass density (g/cm3 ) (at 20◦ C)
0.877
––
1.11
1.29
molar volume (L/mol) (at 20◦ C)
0.089
––
0.101
0.114
enthalpy of formation (kJ/mol)
49.0
0.0
11.0
−18
heat capacity (J/mol K)
136.3
––
150.1
162.4
relative volatility
13
––
3.0
1
enthalpy of vaporization (kJ/mol)
30.72
––
35.19
39
Notes: Normal boiling point is at 1 atm. Mass densities are in the liquid phase at 20◦ C.
Enthalpies of formation are for the liquid phase at 298.15 K and 1 bar from the elements in their
standard states at 298.15 K and 1 bar. Heat capacities are for the liquid phase at 298.15 K and
1 bar. Relative volatilities are with reference to dichlorobenzene (the heaviest species) and are
estimated from data obtained from DISTIL6 using a phase equilibrium model with the NRTL
activity coefficient model for the liquid and ideal vapor phase. Physical property data are taken
from CRC Handbook of Chemistry and Physics.7
Table 2: Stream Table Determined by Shortcut Calculations (kmol/hr)
Stream:
1
2
3
4
5
6
7
Benzene:
52.2
0
219.3
219.3
0
0
0
Chlorine:
0
54.3
0
0
0
0
0
chlorobenzene:
0
0
50.0
0
50.0
50.0
0
dichlorobenzene:
0
0
2.18
0
2.18
0
2.18
6
Table 3: Stream Table Determined by HYSYS (kmol/hr)
Note:
∗
Stream:
1
2
3
4
5
6
7
Benzene:
52.2∗
0
219.0
219.0∗
0
0
0
Chlorine:
0
54.4
0
0
0
0
0
chlorobenzene:
0
0
50.0∗
0
50.0
50.0
0
dichlorobenzene:
0
0
2.18∗
0
2.18
0
2.18
identifies flowrates which were specified as inputs in HYSYS.
Table 4: Equipment Ratings Determined by HYSYS and Shortut Calculations
Item
Shortcut
HYSYS
26.4
29.6
Minimum reflux ratio
0.37
0.32
Vapor Flow Rate (kmol/min)
381
381
Reboiler Heat Duty (mW):
3.7
3.3
Condenser Heat Duty (mW):
3.7
3.3
Minimum reflux ratio
0.52
0.36
Vapor Flow Rate (kmol/min)
102
102
Reboiler Heat Duty (mW):
1.0
1.0
Condenser Heat Duty (mW):
1.0
1.0
Reactor
Reactor volume (m3 ):
Column 1
Column 2
7
Table 5: Equipment Sizes for Nominal and Flexible Process
Item
Nominal
Flexible
Reactor
Reactor volume (m3 ):
26.4
34.3
Column 1
diameter (m)
2.0
2.2
number of trays
45
45
height (m)
29
29
reboiler area (m2 )
82
106
condenser area (m2 )
119
155
Column 2
diameter (m)
0.88
1.00
number of trays
51
51
height (m)
32
32
reboiler area (m2 )
48
63
condenser area (m2 )
24
31
C 6H 6
4
C 6H 5C l
6
C 6H 6
1
2
C l2
CS TR
3
5
C 6 H 4 C l2
7
Figure 1: Process Flow Diagram for Benzene Chlorination
8
E P 4 H10 6 $•yrL
3.5
X
3.0
2.5
2.0
0
100
200 300 400
R B Hkmol•hrL
500
Figure 2: Design-stage optimization of the nominal benzene chlorination process
E P 4 H10 6 $•yrL
3.2
X
3.1
3.0
2.9
100
200 300 400
R B Hkmol•hrL
500
E P 4 H10 6 $•yrL
Figure 3: Operational optimization of the flexible benzene chlorination process
0.6
0.5
0.4
0.3
0.2
0.1
X
0
100
200 300 400
R B Hkmol•hrL
500
Figure 4: Operational optimization of the flexible benzene chlorination process after production
rate decrease.
9