Question: What is the hydrogen supply ratio (alpha) defined as the ratio of hydrogen supplied to stoichiometric requirements in continuous catalytic hydrogenation reactors? Extract numerical values or ranges for typical hydrogen excess ratios used in practice.

Chemical Engineering Science, Vol. 43, No. 8, pp. 2229–2236, 1988.
Printed in Great Britain.
0009-2509/88 $3.00+0.00
Pergamon Press plc
DEVELOPMENT OF CATALYTIC HYDROGENATION REACTORS FOR THE FINE CHEMICALS INDUSTRY
K.R. Westerterp, K.B. van Gelder, H.J. Janssen and M.H. Oyevaar
Chemical Reaction Engineering Laboratories
Chemical Engineering Department
Twente University
P.O. Box 217, 7500 AE Enschede
The Netherlands
Abstract
A survey is given of the problems to be solved before catalytic hydrogenation reactors can be
applied in a multiproduct plant in which selectivity problems are experienced. Some results are
reported on work done on the reaction kinetics of two multistep model reactions and on
mathematical modelling and experimental verification of the models. Since hydrogenation reactions
are often very exothermic, cooling by solvent evaporation has been applied where appropriate.
Sufficient information has been collected and correlated to enable operation of multiproduct
catalytic reactors of the slurry or packed bubble column type; interdependence of operating
variables is so complex that a mathematical model is indispensable.
1. Introduction
In our laboratories cooperation with industry is intensive. It is our job to study the basic
fundamentals for new operations or processes, which will be further elaborated by our partners and
eventually developped on a larger scale into new processes and syntheses. This approach places us
in a position to publish on our results. Since 1984 we are a.o. engaged in developping catalytic
hydrogenation processes for the fine chemicals industry. In this undertaking we have a partner who
has a production program which includes around twenty different hydrogenation reactions of a
variety of raw materials. These hydrogenations are carried out according to the classical Béchamps
process using Fe-powder and a HCL solution. As a consequence large amounts of iron hydroxides and
neutralization salts have to be disposed of, a very costly affair. So it was obvious to look into
possibilities for catalytic hydrogenation.
In catalytic hydrogenation a number of problems can be anticipated. It is too costly to build say
20 different small hydrogenation units if these reactions have to be carried out under different
operating conditions. As a consequence one or two larger units have to be developped for
multiproduct use. Hydrogenation reactions often are highly exothermic - in our case adiabatic
temperature rises of 200-500 K are common - so methods have to be devised to cope with high heat
production rates. In case of producing an intermediate product in a set of consecutive reactions
or of combined reactions selectivity is a problem and operating conditions must be kept under
accurate control. The number of catalysts to be used preferably should be small because of
handling and recovery costs. It would be best to have one single, universal catalyst.
Financial losses should be kept to a minimum. For raw materials this requires short and quick
start-up and stop procedures, for hydrogen it requires low excess of hydrogen or recirculation of
non converted material. For the catalyst it demands a good recovery system and in case of
deactivation simple reactivation procedures. Therefore, a number of problems has to be solved
before catalytic hydrogenations can be introduced on a large scale in a fine chemicals plant.
2. Development objectives
In order to solve the problems at hand we decided to opt for continuous operation of the reactor
and to use an evaporating solvent to remove the reaction heat in the reactor. Miniplants have been
built for catalytic hydrogenations in a slurry reactor, in a packed bubble column and in a
cocurrent trickle flow column reactor. In addition, two separate miniplants have been build for
the study of gas-liquid interfacial areas in several different reactor types, one for pressures in
the range of 0.1-2 MPa and the other one for 0.1-10 MPa. To test the performance of the reactors
experimentally two model hydrogenation reactions have been selected and the reaction rates of the
individual reaction steps have been determined. Mathematical models have been developed to
describe the different reactor types and their behaviour. The models have been tested
experimentally. Interfacial areas are determined by a chemical method using systems of amines and
CO2-gas. In this communication we report briefly on some of the results we have obtained. Some
fifteen papers are being prepared to report more extensively on our results.
3. Model reactions
Having considered many possible complex reactions we decided upon the catalytic hydrogenation of
2,4-dinitrotoluene (DNT) and of 2,4,6-trinitrotoluene (TNT) in methanol and as our model
2229
2230
K. R. WESTERTERP et al.
G12
reactions, and supported palladium as the catalyst. The heat of reaction of the reduction of a
single nitrogroup is around 600 kJ/mole; at reasonable concentrations of TNT adiabatic temperature
rises of 300 to 500 K occur. On theoretical grounds we can expect that the hydrogenation of DNT
proceeds in two parallel paths, each of which consist of two consecutive reactions when
amino-nitro compounds are formed as the sole intermediates. The hydrogenation of TNT involves an
even more complicated scheme.
A large amount of literature has been published in recent years concerning the experimental
investigation and the modelling of the catalytic hydrogenation of several nitrosubstituted
aromatic compounds. Most of the literature refers to the hydrogenation of nitrobenzene. In
general, aromatic nitrocompounds are converted to the amines quantitatively. Commonly Ni and
supported Pd or Pt are used as catalysts. In general the reaction rates reported are zero order in
the nitrocompound and between zero and first order in hydrogen. They can be described with
Langmuir-Hinselwood rate equations. Burge et al. (1980) reported that the hydrogenation of
nitrobenzene, using Raney nickel as a catalyst, did involve the formation of azoxy- and of
azobenzene as stable intermediates. The formation of azoxybenzene as an intermediate has also been
reported by Yao and Emmet (1962). Burge concluded that the reaction mechanism followed a reaction
pathway similar to that proposed by Haber (1898) for the electrochemical reduction of
nitrobenzene. Collins et al. (1982) reported that the reaction pathway of the reduction of
nitrobenzene also obeyed Haber's scheme when nickel boride was used as a catalyst. The maximum
amounts of azoxy- and of azobenzene which builded up during the reaction were found to be
negligable. Bird and Thompson (1980) stated that the reduction of any nitrogroup on an aromatic
ring, using supported palladium as a catalyst, will follow the same pathway as chemical reduction,
but that it is very difficult to terminate the reaction at one of its intermediates. Yucelen
(1984) studied the kinetics of the hydrogenation of 2,6-dinitrotoluene using Pt/C and Pt/Al.
catalysts. The reaction proceeded according to a simple scheme of two consecutive reaction A → B →
C, in which B represents the only intermediate 2-amino-6-nitrotoluene and where C is the
end-product 2,6-diaminotoluene.
11203 as
Although some studies have been published with respect to the catalytic hydrogenation of 2,4-DNT,
practicable models of this reaction and of the hydrogenation of TNT have not yet been developed.
Acres and Cooper (1972) and Bird and Thompson (1980) reported about the accurrence of two
intermediates in half-hydrogenated samples of 2,4-DNT: 2-amino-4-nitrotoluene and 4-amino-2-
nitrotoluene, respectively. Unfortunately they did not give any further information, neither on
the change of the concentrations of the substrate and the products as a function of the course of
the reaction, nor on the analytical technique used. Pawlowski and Kricsfalussy (1981) found a
single step mechanism for the conversion of 2,4-DNT to 2,4-DAT. Acres and Cooper (1972) have found
2-amino-4-nitrotoluene (2-A-4-NT) and 4-amino-2-nitrotoluene (4-A-2-NT) as stable intermediates in
half hydrogenated samples of DNT.
The kinetics of the hydrogenations of DNT and TNT have been investigated in a fully automated
batch slurry reactor, which was operated isothermally and under constant pressure. As a catalyst
5% Pd on active carbon was used. The consumption of hydrogen during reaction was calculated from
the drop in a hydrogen pressure. A set of preliminary experiments has been performed at several
temperatures, agitation rates, catalyst and substrate concentrations to warrant the exclusion of
any mass transfer effects that possibly mask the real reaction rates. For both DNT and TNT sets of
hydrogenation experiments have been carried out at different hydrogen pressures to determine the
effect of the temperature and hydrogen pressure on the overall conversion rates. Routine analysis
were performed with a gas chromatograph which is equipped with a flame ionization detector and
fitted with 6 ft., 1/8 inch column packed with Tenax-GC, 60-80 mesh. So far we have concentrated
mainly on quantitative analyses of reaction mixtures of DNT. The concentrations of the relevant
reaction compounds as a function of time have been determined at reactor temperatures of 308, 323
and 357 K and at a constant hydrogen pressure of 2 MPa. In the samples DNT, 2-A-4-NT, 4-A-2-NT and
DAT have been identified by means of comparison with the retention times of the pure compounds.
Two other products have been identified as 4-hydroxylamino-2-nitro-toluene (2-N-4-HAT) and 2-
amino-4-hydroxylamino-toluene (2-A-4-HAT) by means of GC-MS analysis. Note that the hydroxylamines
are formed at the para-position only and that an ortho-nitrogroup is converted directly to an
aminogroup. The concentrations of 2-A-4-HAT were found to be so low that the pressure of this
compound has been neglected.
A plot of the hydrogen conversion as a function of the reaction time during a typical hydrogena-
tion of DNT at 308 K is presented in Fig. 1. In Fig. 2 the concentrations of the major compounds
are plotted versus the hydrogen conversion. In Fig. 2 also the reaction pathway can be observed.
At first DNT is converted in two parallel reactions into 2-A-4-NT and 2N-4HAT. The consecutive
reactions do not start before almost all of the DNT has been consumed. This is known as molecular
queueing and can be explained in terms of a competitive adsorption of the reacting species on the
catalyst surface. Once all of the DNT has been consumed the 2-N-4-HAT is converted into the amine
4-A-2-NT which subsequently reacts to DAT, while in the parallel pathway simultaneously the
2-A-4-NT is converted to DAT. In Fig. 1 a sharp decrease is shown in the overall reaction rate to
a conversion level of 0.4.
As a consequence of the strong molecular queueing in the initial stage
the reaction exhibits zero order behaviour with respect to DNT. The observed decrease of the
overall reaction rate was caused by rapid deactivation of the catalyst in an induction period.
The kinetic rate data of the hydrogenation of DNT and TNT have been modelled using Langmuir-
Hinshelwood rate equations. The rate of conversion of species i into species j is given by rij
ccat Kij Hi; where OH and i are the relative concentrations on the catalyst surface of hydrogen
and of species i respectively. The experimental results indicate separate active sites for
hydrogen and for the aromatic molecules on the catalyst surface. Assuming dissociative adsorption
of hydrogen on the catalyst Oд and O₁ are given by:
✓(кHPH2/RT)
=
он +(PRT)
and
Qi Xi
0
i
(KACAO ¯ +QjXj
G12
Development of catalytic hydrogenation reactors
2231
KA/C
is
where к is the product of the solubility constant of hydrogen in methanol and the true adsorption
constant, Q₁ is the relative adsorption strength of species i with respect to DNT or TNT,
the adsorption constant, CAO is the initial concentration of DNT or TNT and further Xi
The model for the hydrogenation of DNT is based on the reaction pathway presented in Fig.
was found that the overall hydrogen consumption rate of the hydrogenation of TNT could be
described accurately over the whole temperature range of the experiments by a lumped model of
three consecutive reactions TNT DNHAT DANT→TAT. Unfortunately we lacked sufficient quantitative
data of the liquid phase concentrations of the reaction components to set up a realistic extended
model for the hydrogenation of TNT.
1.00
1.00
4HA-2NT — 4A-2NT
DNT
DAT
2A-4NT
0.80
0.80
0.60
hydrogen
conversion
0.40
0.20
hydrogen pressure: 1.90 MPa
reactor temperature: 308 K
catalyst concentration: 1.203 kg/m³
initial DNT concentration: 0.220 kmole/m
0.60
CiCo
0.40
0.20
0.00
0.00
0
1000
2000
3000
4000
5000 6000
0.00
0.20
0.40
0.60
0.80
1.00
reaction time
hydrogen
conversion
Figure 1:
Hydrogen conversion vs. reaction time for a typical hydrogenation
of DNT in a mechanically agitated three-phase batch reactor.
Figure 2:
Relative concentration C/C vs. hydrogen consumption for a typical
hydrogenation experiment of DNT (reaction conditions see figure 1).
Symbols: DNT, 4HA-2NT, ♦ 4A-2NT, 2A-4NT, □ DAT.
For the parameter estimation of both models we used the RKPES computer program developed by Klaus
and Rippin (Klaus and Rippin, 1979) which applies the maximum likelihood approach. The parameters
have been estimated for each experimental temperature separately. The temperature dependent
relations for the parameters have been derived from the single temperature results. The relations
are summarized in Table 1. The data for DNT comprised the data of the hydrogen conversion as a
function of the reaction time of three different experiments (100 to 120 points) in addition to th
data of the concentrations of the reaction components as a function of the reaction time (20 to 25
sets of 5 concentrations) for each temperature. For TNT only data of the hydrogen conversion as a
function of the reaction time have been used. The t-values of the estimated rate parameters range
up from around 10 to 60 whereas the t-values of the adsorption parameters are poor. This can be
explained by the large difference in the relative adsorption strength of the reacting
components on the catalyst: when a component i is strongly adsorbed and nearly covers the whole
catalyst surface the relative surface concentration i 1 and has thus become insensitive to the
exact
value of the adsorption constant.
Table 1: Parameter values in the rate equations for the hydrogenation of DNT and TNT.
==
substrate and
reaction constants
adsorption parameters
reaction pathway
preexponential
factor
[mole/kg cat s]
activation
temperature
(K)
preexponential
factor
activation
temperature
(K)
DNT
12
*23
1.31 10
5.14 1013
10
7500
K1
10000
2.
k34
2.60 10º
7400
*15
2.13 1011
8800
៩៩៩
0.15
130
[m³/mole]
-3000
vapour
line
condensor
[-]
2500
7.11
[-]
1100
gas out
condensate returned
Q4
0.011
(-]
-320
k≤4
7.17 1013
10200
Qs
0.4
[-]
660
reactant feed
catalyst feed
hydrogen feed
stirred
tank
reactor
product out
TNT
*12
3.3 1011
8800
K₁
-10
3.8 10
1 2 3 4
*23
K34
1.2 1010
[m/mole] -9100
8100
Q2
1600
(-)
1.8 10.10
8100
Q3 1.8 10
(-)
3500
7200
Figure 3:
Outline of a continuous stirred three-phase tank
reactor with an evaporating solvent.
For both systems: K₁– 9.14 10¹8
°exp (10360/T) [m³/mole]
4. Mathematical modelling and experimental confirmation
4.1 Continuous stirred three-phase_slurry_reactor
A series of catalytic hydrogenation experiments has been carried out in a mini-installation with a
three-phase slurry reactor using an evaporating solvent to control the temperature. In Fig. 3 the
basic line-up of the installation is presented. A solution of the reactant, a concentrated slurry
of the catalyst and hydrogen gas are fed to the reactor separately. We assume that the liquid and
the gas phase in the reactor are well mixed and that the gas phase is saturated with solvent
vapour. The reaction products and the catalyst leave the reactor in the product stream. The excess
of non-converted hydrogen gas leaves the reactor via the vapour line to a condensor and carries
away the evaporated solvent, thus withdrawing heat from the reactor. The solvent is condensed and
returned to the reactor: the excess gas leaves the system via a back-pressure control
valve.
At constant reactor pressure and constant hydrogen excess the heat withdrawn from the reactor by
solvent evaporation will rise strongly with increasing temperature because of the exponential
2232
K. R. WESTERTERP et al.
G12
increase of the rate of evaporation. The maximum attainable reactor temperature, the boiling point
of the reaction mixture, can never be reached, because the rate of evaporation would become
infinitely large. The rate of heat production is also affected by the temperature because the
partial pressure of hydrogen is lowered by the increase of the solvent vapour pressure. At a
certain temperature the rate reducing influence of a lower partial pressure of hydrogen will
become larger than the rate enhancing effect of a higher temperature. Consequently at increasing
reactor temperatures the rate of reaction passes through a maximum and then decreases. At the
boiling point the reaction stops completely because the hydrogen cannot dissolve anymore into the
boiling liquid (Westerterp and Crombeen (1983).
A mathematical model of the slurry reactor has been developed which describes the hydrogen
conversion, the rate of heat production (HPR), the total rate of heat withdrawal (HWR) and the
liquid phase composition as a function of the reactor temperature.
The material balances of the reaction components are solved simultaneously using the kinetic model
of the reaction system. Thus the overall conversion rate of hydrogen gas and the HPR can be
determined. Using the lumped reaction model, for the hydrogenations of TNT we write:
v,1°C
$x,1 (CAO CA
and CD
=
CAO
(CA+
+
VR, 1 AB
св °C)
==
*v,1ºB VR,1 (rAB rBC)
VR,1 (BC r CD)
=
с catki H1 is the rate
Where A, B, C and D represent TNT, DNHAT, DANT and TAT respectively rij
of the reaction of i into j. The material balance for the hydrogen is given by:
*H₂out = "H₂in *v,1 H2A0H2
0 v,1 H₂ Ao represents the amount of moles of hydrogen required for complete conversion of the TNT
in the feed to TAT, H₂ is the hydrogen conversion. * H2in is related to the reactant feed by
*H₂in* *v,1ª¹H₂˚Ao' in which a is defined as the amount of moles of H₂ fed to the reactor over the
amount required for complete conversion. Now we can write:
*H₂out = *v,1°H2ºAo(a
-
5H2
>
The HPR is calculated on the assumption that the heat liberated by the reaction is directly
proportional to the rate of hydrogen conversion in the reactor:
HPR =
v,1ºAoH2 (HR)
For the HWR three contributions have been distinguished:
1) the heat absorbed by the cold feed streams of the reactant solution, the catalyst slurry and
the hydrogen gas:
v, feed (pCp) 1 (T-Tfeed)
+
*v, cat (pCp) 1 (T-Tcat)
+
с
,1°VH2Ao (MCp) H2 (T-TH2)
HWR ef
2) the heat absorbed by evaporation of the solvent. We assume that the gas and the liquid in the
reactor are at equilibrium. The vapour stream that leaves the reactor has the same composition
as the gas-phase in the reactor and consists of hydrogen, methanol and reaction water. It is
related to H₂out by:
HWR
=
ev
vap
*H₂out р (pgxs+ Ph*w)
vapW
The denominator of this equation represents the hydrogen pressure in the reactor.
3) the heat absorbed by the condensate returned to the reactor. We assume complete condensation of
the vapour in the condensor, so we write:
HWR
=
cond *H₂out
PSXs (MCD)s+ PWW (MCD)W (T-T,
Р (poxs+ pwxw)
cond)
In Fig. 4 the results of model calculations of the hydrogenation of TNT are presented for a
typical set of operating conditions. In Fig. 4a the HPR and HWR and in Fig. 4b the concentrations
1,5
1,0
Relative
1,0
HPR and HWR
|
0,5
HPR
1
0,8
TNT
DN-HAT
DA-NT
G/C,TNT
0,6
0,4
TAT
0,2
0,0
30
0,0
50
70
reactor temperature
90
110
30
50
70
90
110
reactor temperature
с
Figure 4a:
Results of calculations with the reactor model.
Dimensionless heat production rate (HPR) and heat
withdrawal rate (HWR) as a function of the reactor
temperature at constant residence time of the liquid.
Constant catalyst food and gas feed temperature of 20°C.
Feed temperature: a: 20°C, b: 30°C, c: 50ºC, &: 70ºC.
Figure 4b:
Results of calculations with the reactor model.
Dimensionless concentrations G /CTNT,o in the
reactor product as a function of the reactor temperature
at the same operating operating conditions as in
figure 4b case b. The dashed lines indicate the lower
and upper stable operating point, respectivily.
G12
Development of catalytic hydrogenation reactors
2233
of the components in the reactor product are plotted versus the reactor temperature at constant
residence time of the liquid in the reactor. The HPR and HWR have been normalized by dividing them
by the maximum HPR at complete conversion of the TNT. The concentrations have been normalized by
division by the feed concentration CAO. Because the calculations have been based on the lumped
kinetic model of TNT the concentrations plotted in Fig. 4b do not represent the true product
distribution in the reactor but only give an impression of how the concentrations of the reaction
components change as a function of the reactor temperature.
Because the kinetic model does describe the hydrogen consumption rate very well the heat effects
calculated with the reactor model are realistic. Curves a, b, c and d in fig. 4a represent the HWR
at four different feed temperatures. The reactor works at a stationary operating point when the
stability criteria HWR-HPR and SHWR/8T > SHPR/ST are fullfilled. In Fig. 4a multiplicity occurs at
the feed temperatures of the a and b curves. In Fig. 4b the dashed lines indicate the lower and
upper stable operating point of case b. The relative concentrations in the reactor outlet stream a
the lower stationary point are 0.42 TNT, 0.55 DNHAT and around 0.03 DANT and in the upper
stationary point are 0.04 TNT, 0.10 DNHAT, 0.51 DANT and 0.35 TAT. Multiplicity can be avoided by
using a slightly higher feed then used for the calculations of case b. The most important
parameters which control the behaviour of the reactor are the total pressure in the reactor, the
residence time of the liquid phase in the reactor and the temperature of the reactant feed.
In total of 22 hydrogenation of TNT experiments have been carried out in the mini-installation. By
varying the operation conditions a temperature range of around 35 to 100°C has been covered. In
Fig. 5 the experimental results are compared with the results of model calculations with respect
to the hydrogen conversion and the temperature at the operating point. The agreement is
satisfactory, especially when realising that small errors in the calculation of the HPR or HWR may
result in a drastic change of the predicted steady state temperature.
1.00
0.80
0.60
I
"
hydrogen
conversion
0.40
I
DIO
0.20
model calculations
experimental results
0.00
30
50
70
90
reactor temperature
C
110
Figure 5: A comparison of experimentally determined hydrogen conversions
and reactor temperatures and those predicted by our model.
4.2 The three-phase_packed_bed_bubble_column_reactor
A model comprising plug flow and axial dispersion of mass and heat is used to describe the
behaviour of the packed bed reactor. The derived model equations are based on the following
assumptions: 1. the reactor is adiabatic and isobaric, 2. the gas phase is in equilibrium with the
liquid phase when it enters the catalyst bed, 3. the gas and the liquid feed temperature are
equal, 4. there are no mass transfer limitations, 5. the temperature in all three phases is
uniform in a plane perpendicular to the direction of flow, 6. there are no concentration profiles
in the radial direction, 7. the reactants are not volatile, 8. there is no heat conduction in the
axial direction, 9. the gas phase is ideal, 10. the contribution of the reactants to the specific
heat and the molar weight of the liquid phase can be neglected, 11. the available catalyst area is
fully wetted and utilised, 12. a mass transfer rate equation is used to describe the evaporation
of the solvent. The partial pressure of the solvent at the gas-liquid interphase is equal to the
vapour pressure of the solvent.
The gas-liquid equilibrium at the entrance of the catalyst bed is achieved by assuming that the
gas and the liquid feed flow through an adiabatic mixer in which this equilibrium is reached
before entering the catalyst bed. Since some evaporation will occur in the mixer, the temperature
at the entrance of the catalyst bed will be lower then the feed temperature.
The dimensionless equations based on the above assumptions are:
1
* for the reactants
-
dz
dx.
-
do
+ 8
* for hydrogen in the gas phase YH dz
* for the solvent in the gas phase
* for the liquid phase az +
* enthalpy balance # d²
i
Y
ᏓᎾ
St' P
G
+ Da
= 0
n R prod, i
a H
dz
Da n
α
Rprod,
0
d Y
sol
+
H 1
sol
St P (
dz
sol)
ΨΘ
0
0
Фо
of
(S01 ✗sol) ½
-
Y
+40
+
+
Pen az
CPGMGG)
= о
* from the definition of mole fractions X sol
The boundary conditions for these equations are:
de
-Da
n Σ(R, ATad,j)
St' PATev,sol
=
Σ X j
and
хдо
1
de
* at Z = 0
(CPGMG®GB
+
°L) (e
- 0 +)
+
Z=O
Z=O
Pen
dz
= °
=
1
2234
K. R. WESTERTERP et al.
-pem
X
and
i, Z-O-
==== 1 -
Y sol
$L
=
Φ
L, Z=0°
LO
G
G,Z=O
d X.
| z-o-
-
Y sol
=
Y
sol,
Pem
Z=O¯
~1,Z=0+
YH
de
* at Z=1
dx i
= 0
G12
The conditions at Z=0¯ follow from the balances around the mixer. The dimensionless groups and
variables are defined in the notation. The dimensionless temperature 0 is defined in accordance to
the definition used by Westerterp and Crombeen (1983).
The hydrogenation of TNT to TAT was chosen as a model reaction in our cocurrent upflow packed bed
reactor. The kinetic rate equations presented in section 3 were used in the model calculations.
The catalyst used in the reactor is a shell catalyst of 0.08wt% Pd on 4.2*4.2 mm porous
cilindrical pellets. It was assumed that the carrier does not influence the kinetic parameters. An
allowance is made however, for a difference in the activity of the catalysts. The term Ccat is
replaced by n no Pp. in which pp is the packing density, a scaling factor and In the parameter
describing the catalyst activity.
ת
Two important operating parameters are the hydrogen supply ratio a and the reactor pressure. The
hydrogen supply ratio a is defined as the ratio of the number of moles of H₂ fed to the reactor to
the number of moles of reactant fed to the reactor. If n₁ is defined as the number of moles of the
H2 needed for complete conversion of all reactant into the desired product i, we can define a
supply ratio a₁ as: di a/n. With this definition a value of a; 1 means that the exact amount
of H2 is fed to the reactor necessary to convert all reactant into product i. For industrial
purposes it can be advantageous to keep a₁, if i is the desired product, close to 1 to avoid
hydrogen losses or high recycle ratios for unconverted hydrogen. On the other hand, if the value
of
=
is chosen close to 1 the hydrogen partial pressure will drop to very low values because
practically all hydrogen is consumed. Consequently the concentration of hydrogen at the catalyst
surface will decrease, resulting in lower reaction rates and the need for larger reactors.
Furthermore, the value of a determines how much hydrogen is present and thus how much solvent
evaporates. Therefore, a also has a strong influence on the amount of heat that is absorbed by
evaporation and thus on the temperature level in the reactor.
The reactor pressure determines the boiling point of the solvent and thus the maximum temperature
that can be reached in the reactor.
Typical results from a model run are given in Fig. 6a-d. Some important effects can be noted from
these figures:
because more then one mole of solvent evaporates for each mole of H₂ consumed, the gas flow rate
increases.
- the mole fraction of H2 drops to low values near the end of the catalyst bed due to the
evaporation of the solvent.
The influence of CTAT and the reactor pressure on the fraction of the reaction heat that is
absorbed and removed by evaporation of the solvent is given in Fig. 7.
A comparison between the temperature and conversion profiles measured during an experimental run
and calculated with the model are given in Fig. 8 a and 8b. The parameters that were varied to
obtain a best fit the model and the experiment are Pe
From this figure we
conclude that the between well describe the experimentally determined temperature and unversion
profiles. Since the kinetic data available at the present time do not allow calculation of the
composition of the liquid phase, no conclusions can be drawn concerning the selectivity to the
reaction intermediates.
1.0
↑
0.8
0.6
5
[-]
0.4
0.2
3
vapour
2
Ga
[-]
1
Liquid
6c
0.0
0
0.0
0.2
0.4
0.6
0.6
1.0
0.0
0.2
dimensionless length (-)
0.4
dimensionless length [-] -
0.6
0.8
1.0
110
100
T
90
[°C]
80
1.0
0.8
0.6
Ун
ob
[-]
0.4
0.2
6d
0.0
70
0.0
0.2
0.4
0.6
0.8
1.0
0.0
0.2
0.4
0.6
0.8
1.0
dimensionless length [-]
dimensionless length [-]
Figure 6: Calculated profile for the packed bed reactor model.
Model parameters: P10, Per-15, P-500 kPa, OTAT-2.4, food temperature = 50 °C, adiabatic temperature rise-170 °C, Da-5.
a. conversion, b. temperature, c. dimensionless flows, and d. mole fraction of H₂ in the gas phase
G12
[-]
100
P-0.2 MP
90
P-0.4 MPa.
80
|
70
P=0.6 MPa
60
°
1
2
3
$
α [-]
Development of catalytic hydrogenation reactors
110
Calculated
■ Experimental
100+
T
(℃)
90
1.0
0.8
0.6
H}
0.4
2235
0.2
Calculated
Experimental
BO
0.0
0.0
0.2
0.4
0.6
0.8
1.0
0.0
0.2
0.4
0.6
0.8
1.0
Z [-]
Z (-)
Figure 7: Percentage of the reaction heat (at 100% conversion)
that is used for evaporation of the solvent. Model parameters are:
P-10, Po-1.5, feed temperature - 55 °C and Du - 3.5.
Figure 8: Comparison between experimental data and calculated profiles. Model parameters are:
-2
111
■ 5.0, Pe -0.9 and St' - 760.
h
TNT,0
Da = 6.95, α TAT = 2.4, y = 5.38, n = 0.365, Pe
Experimental data: – 1.06·10 ´ m/hr, ♦ = 1.08 Nm/hr, T-50 °C, x, -0.00873
adiabatic temperature rise - 170°C, P - 490 kPa and reactor length - 1 m.
a. temperature profile, b. conversion profile.
5.
Discussion and conclusions
In a bird's eye view we presented some results of our program for the development for catalytic
hydrogenation reactors for multi-product plants. We focussed on the kinetics of our model
reactions, the modelling of our reactors and on the experimental verification of our models.
We showed both theoretically and experimentally - that for every exothermic reactions heat
can easily be achieved by evaporation of a solvent. It is now possible to install the
cooling area outside of the reactor. Temperature control can be easily achieved by adjusting the
reactor pressure.
removal
A disadvantage of the evaporation of the solvent is the subsequent decrease of the partial
pressure of hydrogen. This may result in a decrease of the rate of reaction. The decrease in the
partial pressure of hydrogen can be compensated for by increasing the reactor pressure, by using a
larger excess of hydrogen or by employing larger interfacial areas.
The interdependence of all the operating variables of the reactors is so complex, that a catalytic
hydrogenation reactor in a multi-product plant can not be run without the use of a mathematical
model that describes its behaviour. The good agreement observed between our models and the
experimental data gives us the confidence that our models will be adequate for this purpose.
The most difficult aspect is the acquisition of sufficiently reliable kinetic data. For fine
chemicals the tremendous effort required to obtain these data with present methods is prohibitive.
New methods have to be developed to this end. This will be a subject of further study, especially
for the hydrogenation of TNT.
Acknowledgement The investigations were partly supported by the Dutch Technology Foundation STW
and partly by Andeno BV, Venlo, the Netherlands.
References
1. Acres, G.J.K. and Cooper, B.J., 1972, J. Appl. Chem. Biotechnol., 22, 769-785.
Bird, A.J. and Thompson, D.T., 1980, Academic Press, New York.
2.
3.
Burge, H.D., Collins, D.J. and Burtron, H.D., 1980, Ind. Engng. Chem.
389-391.
Prod. Res. Dev., 19,
4.
5.
6.
7.
8.
9.
Collins, D.J., Smith, A.D. and Burtron, B.H., 1982, Ind. Engng. Chem. Prod. Res. Dev., 21
279-281.
Haber, F., 1898, Z. Elektrochem., 22, 506.
Kan, Kin-Mun and P.F. Greenfield, 1983, A. I.Ch.E.J. 29, 123-132.
Klaus, R.A. and Rippin, D.W.T., 1979, Comp. & Chem. Engng., 3, 105, see also Klaus, R.A.,
1981, Thesis E.T.H. Zürich.
Michelsen M.L. and K. Østergaard, 1970, Chem. Eng. Sci. 25, 583-592.
Pawlowski, J. and Kricsfalussy, Z., 1981, Chem.-Ing.-Techn., 53, 652-654.
10. Yao, H.C. and Emmet, P.H., 1962, J. Am. Chem. Soc., 84, 1086-1091.
11. Yücelen, F., 1984, Thesis E.T.H. Zürich.
12. Westerterp K.R. and Crombeen P.R.J.J., 1983, Chem. Eng. Sci., 1331-1340.
13. Westerterp K.R., van Swaaij W.P.M. and Beenackers A.A.C.M., 1984, John Wiley & Sons,
Chichester, UK.
Notation
2A-4NT
2A-4HAT
2N-4HAT
4A-2NT
ag
Во
C
C
cat
ср
2-Amino-4-Nitrotoluene
2-Amino-4-Hydroxylaminotoluene
2-Nitro-4-Hydroxylaminotoluene
4-Amino-2-Nitrotoluene
specific gas/liquid interfacial area
Bodenstein number
concentration, [mol m³]
concentration of catalyst, [kScatm¯³]
specific heat
dimensionless specific heat of the gas
axial dispersion coefficient
[1/m]
[kJ/kg K]
phase defined as Eypi/p
[m/s]
2236
K. R. WESTERTERP et al.
G12
Da
DANT
DNHAT
DNT
Damköhler number defined as no Pb k₁,S L/OLO XA,O
diamino nitro tolue DAT
dinitro hydroxylamine toluene
2,4-dinitrotoluene
2,4-diaminotoluene
HPR
Heat Production Rate, W
HWR
Heat Withdrawal Rate, [w]
k
k t
K
L
M
MG
ni
P
P1
PR
P
mass transfer coefficient, [ms]
reaction rate constant, [mole s 1kg cat ¯1]
adsorption constant, (m³ mol]
reactor length
molar mass
[m]
[kg/kmol]
dimensionless molar mass of the gas phase defined as Ey, M./MH2.
number of moles of hydrogen needed for complete conversion of 1 mole of reactant to
component i
dimensionless pressure defined as P/Pg
vapour pressure of pure species i,
total pressure
Standard Pressure
[Pa]
[N/m²]
Peclet number for heat dispersion
Θ
r
1
R.
dimensionless reaction rate: ri/pbk 1, s.
prod,i
Peclet number for mass dispersion
relative adsorption strength,
rate of conversion, [mole s¨¹kg,
dimensionless production rate of component i
modified Stanton number (also called NTU)
gat 1]
T
TÅT
TNT
Vr,1
temperature
[K]
standard temperature
2,4,6-triaminotoluene
2,4,6-trinitrotoluene
liquid volume in the reactor, [m³]
liquid mole fraction of species i
dimensionless mole fraction of component i: ×
mole fraction of component i in the gas phase
dimensionless axial coordinate = h/L
A,O
α i
B
Υ
ΔΗ
ev
AHri
ΔΗ
AT.
AT
€ G
vap
ad,i
ev,i
5H2"
5
hydrogen supply ratio: ratio of the number of moles of hydrogen fed to the reactor per unit
of time to the number of moles of TNT fed to the reactor per unit of time
hydrogen supply ratio based on component 1
ratio of heat capacities of the gas and the liquid phase in the feed
molar ratio of methanol to hydrogen in the feed
enthalpy of evaporation
reaction enthalpy
heat of evaporation, (kJ mole¯1]
[kJ/mole]
[kJ/mole of component i converted]
dimensioness adiabatic temperature rise for reaction i
dimensionless heat effect of evaporation for component i
hold up of the gas phase
[m³gas/m³reactor]
hydrogen conversion, defined as the amount of hydrogen consumed by the reaction over the
amount required for complete conversion
catalyst activity
scaling factor from the Pd on Carbon slurry catalyst to the Pd on Alumina catalyst used in
the packed bed reactor: 1.28 * 10°
relative surface concentration
effective axial thermal conductivity
stoichiometric coefficient of hydrogen of the overall reaction
n
По
Ө
0
dimensionless temperature: -(E₁/RT_) (1
T/T
[W/m²K]
λeff
[kg/m³]
Ρ
density
[kgcat/m³reactor]
Pb
specific catalyst mass
Φ
feed rate of hydrogen gås, (mole s¯1]
molar flow rate per unit of cross sectional
dimensionless flow: $100
area
of the reactor [kmol/m²s]
8
H₂in
*H₂out
flow rate of the hydrogen off-gas, [mole s¯¹]
total liquid feed rate to the reactor, [m³s-1]
sub and super-scripts
ax
in axial direction
sol
b
refering to catalyst bed
G
solvent
gas
cat
catalyst, catalyst feed stream
H
cf
cold feed streams
L
hydrogen
liquid
cond
condensate returned
о
ev
evaporation of the solvent
S
feed
reagent feed stream
S
feed conditions
refering to standard conditions
solvent, methanol
1
liquid
W
reaction water
O
initial

Two important operating parameters are the hydrogen supply ratio a and the reactor pressure. The hydrogen supply ratio a is defined as the ratio of the number of moles of H₂ fed to the reactor to the number of moles of reactant fed to the reactor. If n₁ is defined as the number of moles of the H2 needed for complete conversion of all reactant into the desired product i, we can define a supply ratio a₁ as: di a/n. With this definition a value of a; 1 means that the exact amount of H2 is fed to the reactor necessary to convert all reactant into product i. For industrial purposes it can be advantageous to keep a₁, if i is the desired product, close to 1 to avoid hydrogen losses or high recycle ratios for unconverted hydrogen.

Model parameters: P10, Per-15, P-500 kPa, OTAT-2.4, food temperature = 50 °C, adiabatic temperature rise-170 °C, Da-5.

hydrogen supply ratio: ratio of the number of moles of hydrogen fed to the reactor per unit of time to the number of moles of TNT fed to the reactor per unit of time