Continuous hydrogenation of 2-butyne-1,4-diol to 2-butene- and butane-1,4-diols

Article abstract of DOI:10.1021/op050216r

Continuous catalytic hydrogenation of 2-butyne-1,4-diol (B₃D) was carried out in a fixed-bed reactor over 1% Pt/CaCO₃ catalyst to give 2-butene-1,4-diol (B₂D) and butane-1,4-diol (B₁D) without formation of any other side products. In case of continuous hydrogenation, higher selectivity (66%) to B₂D could be obtained and the selectivity pattern was completely different from that found in case of batch slurry operation in which B₁D selectivity was very much higher (83%) than the B₂D selectivity (17%). Another interesting feature was that by varying the contact time, the selectivity to both B₂D as well as B₁D could be varied over a wide range which is an attractive option to obtain the desired products mix of B₁D and B₂D, depending on the fluctuation in the market demand. Further, a mathematical model for reactor performance was also developed on the basis of the kinetic data obtained previously in a batch slurry reactor. The predicted values of conversion, selectivity, and rate of hydrogenation were found to agree with the experimental data over a wide range of conditions.

Full text of DOI:10.1021/op050216r

Organic Process Research & Development 2006, 10, 278−284
Continuous Hydrogenation of 2-Butyne-1,4-diol to 2-Butene- and
Butane-1,4-diols
C. V. Rode,* P. R. Tayade, J. M. Nadgeri, R. Jaganathan, and R. V. Chaudhari
Homogeneous Catalysis DiVision, National Chemical Laboratory, Dr. Homi Bhabha Road, Pune - 411008, India
Abstract:
give 60-70% selectivity to B
2
D while remaining is a mixture
Continuous catalytic hydrogenation of 2-butyne-1,4-diol (B
3
D)
of saturated diol (B D) along with other side products such
1
was carried out in a fixed-bed reactor over 1% Pt/CaCO
catalyst to give 2-butene-1,4-diol (B D) and butane-1,4-diol
D) without formation of any other side products. In case of
continuous hydrogenation, higher selectivity (66%) to B D could
be obtained and the selectivity pattern was completely different
from that found in case of batch slurry operation in which B
selectivity was very much higher (83%) than the B D selectivity
17%). Another interesting feature was that by varying the
contact time, the selectivity to both B D as well as B D could
be varied over a wide range which is an attractive option to
obtain the desired products mix of B D and B D, depending
on the fluctuation in the market demand. Further, a math-
ematical model for reactor performance was also developed on
the basis of the kinetic data obtained previously in a batch
slurry reactor. The predicted values of conversion, selectivity,
and rate of hydrogenation were found to agree with the
experimental data over a wide range of conditions.
3
as γ-hydroxy butyraldehyde, n-butyraldehyde, n-butanol,
2
crotyl alcohol, and acetal due to the double bond migration
,5
(
B
1
and hydrogenolysis of B
2
D in the presence of a catalyst.4
2
It is clear from the literature that, for butyne diol hydrogena-
tion, a monometallic catalyst system gives saturated diol
1
D
(B
whereas a combination of two or more metals or the presence
of organic/inorganic bases gives enhanced selectivity to B D;
1
D) as major product along with other side products,
2
(
2
2
1
however, such catalyst systems lack consistency in activity
in subsequent catalyst reuse, and moreover, such processes
require the complete removal of the additives for obtaining
highest purity of the product for its end use in the fine
chemical or pharmaceutical sector. Also, previous work on
butynediol hydrogenation has mainly been carried out in a
batch slurry reactor using finely powdered catalyst in which
1
2
the selectivity ratio of B
both B D and B D are large-scale commercial products, it
would be most desirable to have a continuous hydrogenation
of B D to give either B D or B D selectively or a desired
mixture of B D and B D. As per the market demand, this
continuous operation for the hydrogenation of B D by merely
changing the operating conditions can give a desired mixture
of B D and B D (Figure 8) for the same catalyst. Hence,
objectives of this work were (i) to investigate the activity
and selectivity of 1%Pt/CaCO catalyst for continuous
2 1
D to B D is normally constant. Since
2
1
3
2
1
Introduction
2
1
Hydrogenation B
3
D in the presence of a catalyst is an
3
industrially important reaction for the manufacture of B
2
D
1
and B
the manufacture of endosulfan and vitamins A and B
whereas B D has a wide range of applications in the polymer
1 2
D. The olefinic diol, B D, is a starting material for
2
1
6
,
1
3
industry and as a raw material for the manufacture of
tetrahydrofuran.1 The earlier processes described Ni- and
Cu-based catalysts for butyne diol hydrogenation under
hydrogenation of butyne diol in a fixed-bed reactor, (ii) to
study the effect of various reaction parameters on the
conversion and selectivity behaviour for the continuous
-3
2
severe operating conditions (15-30 MPa H pressure and
3
hydrogenation of B D, and (iii) to develop a reactor model
4
up to 433 K temperature). The noble metals such as
palladium, ruthenium alone or in combination with other
metals such as zinc, lead, cadmium, copper, and/or organic
amines also were used as catalyst systems to improve
selectivity (ratio of concentration of desired product formed
to the concentration of substrate consumed) to the intermedi-
that provides reasonably accurate conversion and selectivity
predictions for various reactor inlet conditions. For this
purpose, hydrogenation of butyne diol was carried out in a
tubular reactor (30 g capacity); we found that the selectivity
pattern obtained in a fixed-bed reactor was different from
that obtained in the slurry reactor at similar conversion levels
1,5-8
2
ate, B D.
A monometallic Pd/C catalyst was reported to
1 2
and that the selectivity ratio of B D to B D could be altered
*
To whom correspondence should be addressed. Fax: +91 20 25893260.
by varying the H pressure, temperature, liquid and and gas
2
E-mail: cv.rode@ncl.res.in.
flow rate conditions at the reactor inlet. The predictions
obtained by the proposed reactor model were compared with
the experimental data and were found to agree well over a
wide range of operating conditions.
(
1) Winterbottom, J. M.; Marwan, H.; Viladevall, J.; Sharma, S.; Raymashasay,
S.; Heterogeneous catalysis and fine chemicals IV. In Blaser, H. V., Baiker,
A., Prins, R., Eds.; Stud. Surf. Sci. Catal. 1997, 108, 59.
2) Chaudhari, R. V. In Proceedings of the Indo- German Workshop on High-
Pressure Technology Engineering; Chaudhari, R. V., Hofmann H., Eds.;
Forschungszentrum Julich GmbH: Germany, 1993; p 197.
(
(
(
(
3) Telkar, M. M.; Rode, C. V.; Rane, V. H.; Jagannthan, R.; Chaudhari, R. V.
Appl. Catal., A 2001, 216, 13.
4) Rylander, P. N. In Hydrogenation Methods; Katrizky, A. R., Methcohn,
O., Rees, C. W., Eds.; Academic Press: New York, 1985.
5) Bond, G. C.; Webb, G.; Winterbottom, J. M. J. Catal. 1962, 1, 74.
Experimental Section
Materials. B D of >99.5% purity was obtained from
3
E.Merck India, Ltd., and 10% aqueous solution was used.
(
6) Rosso, R.; Mazzocchia, C.; Gronchi, P.; Centola, P. Appl. Catal. 1984, 9, 269.
(8) Bollger, G.; Boer, W.; Wache, H.; Gratze, H.; Koerning, W.; Ger. Pat.
2451929, 1976.
(7) Fukuda, T.; Kusama, T. Bull. Chem. Soc. Jpn. 1958, 31, 339.
2
78
•
Vol. 10, No. 2, 2006 / Organic Process Research & Development
10.1021/op050216r CCC: $33.50 © 2006 American Chemical Society
Published on Web 02/16/2006

Table 1. Range of operating conditions
catalyst
initial concentration of B
solvent
1% Pt/CaCO
3
1.16 kmol/m
water
3
3
D
H
2
pressure
liquid velocity, U
gas velocity, U
10-70 bar
l
10-60 mL/h
20-80 NL/h
0.015 m
0.35 m
g
reactor diameter (i.d.)
total reactor length
catalyst packing length
bed voidage
0.25 m
0.7
particle diameter, d
p
0.004 m
1685 kg/m
0.1050
6.2
3
density of the catalyst, F
porosity of the catalyst, ꢀ
tortosity of the catalyst, τ
p
3
Scheme 1. Hydrogenation of B D
Figure 1. A schematic of the trickle-bed reactor setup.
Hydrogen gas with >99.98% purity was supplied by M/s
Indian Oxygen Ltd., Mumbai, and was used directly from
the cylinder. For the preparation of 1%Pt/CaCO
3
catalyst,
the required quantity of PtCl was dissolved in minimum
4
of time and were analyzed by gas chromatography (HP 6890
series) o.d. 0.32 m, helium gas as carrier, FID detector,
operating temp 80-190 °C using HP-FFAP capillary column.
The liquid velocity was changed after confirming the
steady-state achievement. Following this procedure, the
experiments were carried out at different inlet conditions of
liquid flow rate, gas flow rate. The reactor was operated in
the temperature and pressure ranges of 60-100 °C and 10-
amount of water; if necessary, a small quantity of dilute HCl
was added to ensure the complete dissolution of the precur-
sor. Under stirring, the slurry of support prepared in water
was added to the above solution, and temperature was
maintained at 353 K. After 1 h, formaldehyde was added
under stirring. Then the reaction mixture was cooled, filtered
to obtain the catalyst, which was then dried at room
temperature under vacuum. The powdered catalyst was then
pelletized in the form of pellets of 4-mm diameter.
7
0 bar respectively. The steady state performance of the
reactor was observed by analysis of the reactant and products
in the exit stream.
Fixed-Bed Bench Scale Reactor Setup. All the hydro-
genation experiments were carried out in a bench scale, high-
pressure, fixed-bed reactor supplied by M/s Geomechanique,
France. A schematic of the reactor setup is shown in Figure
Results and Discussion
1
1
. It consisted of a stainless steel tube of 0.35 m length and
.5 × 10 m inner diameter that was heated by two tubular
The main objectives of this work were (i) to investigate
-
2
the activity and selectivity behaviour of 1%Pt/CaCO
3
catalyst
furnaces whose zones (TIC1 and TIC2) were independently
controlled at the desired bed temperature. The reactor was
provided with two thermocouples [Chromel-Alume thermo-
couples (type K)] to measure the temperature at two different
points. The reactor was equipped with mass flow controllers,
pressure indicator, and controller (PIC) devices. A storage
tank was connected to the metering pump through a
volumetric buret to measure the liquid flow rate. The pump
for continuous hydrogenation of B D in a fixed-bed reactor,
3
(ii) to study the effect of various reaction parameters on the
conversion and selectivity behaviour for the continuous
hydrogenation of B D, and (iii) to develop a reactor model
3
for predicting the conversion and selectivity under various
inlet conditions. For this purpose, the effects of temperature,
H
2
pressure, liquid and gas flow rates, catalyst loading on
the conversion of B D and the product selectivities were
3
-
4
3
had a maximum capacity of 3 × 10 m /h under a pressure
of 10 MPa. The other end of the reactor was connected to a
gas-liquid separator through a condenser.
investigated. The range of reaction conditions is given in
Table 1. In all experiments, the concentrations of the reactants
and products in the reactor effluent were determined from
which the conversion and selectivity were calculated. The
material balance as per the reaction Scheme 1, was found to
be in the range of 98-100% in all the experiments.
Reproducibility of the rate, conversion, and selectivity
measurements was found to be within 3-4% error as
indicated by a few repeated experiments. The results are
discussed below.
The experiments were carried out over 10 and 20 g of
catalyst in the form of pellets having 4 mm diameter. The
section 5 cm above and 5 cm below the catalyst bed was
packed with inert packing (carborundum), thus providing the
catalyst bed depth of ∼0.25 m. The reactor was flushed
2
thoroughly with H at room temperature before the start of
the actual experiment. After attaining the desired temperature
the reactor was pressurized with H . The liquid feed was
switched on” after the reactor has reached the operating
2
For 1%Pt/CaCO
3
catalyst, the products of B
3
D hydroge-
“
nation were B D and B
2
1
D as per the reaction Scheme 1.
pressure and kept there for 1 h to obtain the constant liquid
flow rate. Liquid samples were withdrawn at regular intervals
Table 2 shows the variation of conversion and selectivity as
a function of contact time defined as a ratio of catalyst
Vol. 10, No. 2, 2006 / Organic Process Research & Development
•
279

Table 2. Activity and selectivity pattern for butyne diol
hydrogenation in a fixed-bed reactor
selectivity %
W/F
(h)
conversion
(%)
sr. no.
B
1
D
2
B D
1
2
3
4
5
0.33
0.50
0.66
1.11
2
14
30
40
55
78
34
42
49
59
73
66
58
51
41
27
Figure 3. Effect of liquid flow rate on conversion. Reaction
conditions: pH2, 20 bar; H flow rate, 20 L/h; temperature, 100
C; catalyst weight, 10 g.
2
°
Figure 2. Time on stream catalyst activity profile. Reaction
conditions: pressure, 20 bar; gas flow rate, 20 L/h; liquid flow
rate, 20 mL/h; catalyst weight, 20 g.
loading to the liquid flow rate (W/F, h). Obviously, the
conversion of B
3
D increased with increase in contact time.
The selectivity pattern was completely different from that
found in case of batch slurry operation in which B
1
D
D
selectivity was very much higher (83%) than the B
2
9
selectivity (17%). As can be seen from Table 2, varying
the contact time at different conversion levels could vary
the selectivity to both B
This is an attractive option for a commercial operation where
the desired products mix of B D and B D can be achieved,
2 1
D as well as B D over a wide range.
Figure 4. Effect of liquid flow rate on conversion. Reaction
conditions: pressure, 20 bar; H flow rate, 20 L/h; temperature,
100 °C; catalyst weight, 20 g.
2
1
2
depending on the fluctuation in the market demand.
Some preliminary experiments showed that the activity
was increased, the residence time of B
reduced and less time was available for the intimate contact
of H with liquid B D and with the catalyst pellets; thus,
3
D in the reactor was
of 1%Pt/CaCO
for >100 h in the temperature range of 333-373 K (Figure
), confirming the uniform activity for the duration of all
3 3
for the hydrogenation of B D was constant
2
3
proper diffusion was not possible. Another reason is that,
although the increase in the liquid flow rate could wet more
surface area of catalyst pellets and there was an increase in
the gas-liquid and liquid-solid mass transfer coefficient,
at lower liquid velocity, catalyst particles were partially
wetted; under these conditions the rate would increase due
to direct transfer of the gas-phase reactant to the catalyst
surface (already wetted internally due to capillary forces).
Hence, with an increase in the liquid flow rate, an increase
in the wetted fraction was expected to retard the rate of
2
experiments carried out in this work. Therefore, all of the
data collected for the purpose of reactor modeling and reactor
performance studies were under steady-state conditions.
Effect of Liquid Flow Rate. The effect of liquid flow
rate on conversion of B
0 mL/h, under constant temperature, H
flow rate conditions, and the results are shown in Figures 3
and 4. The conversion of B D decreased almost proportion-
ately (from 70 to 12%) with increase in liquid flow rate from
0 to 60 mL/h (see Figure 3). One of the main reasons for
decrease in the conversion is that, as the liquid flow rate
3
D was studied in the range of 10-
6
2
pressure, and gas
3
1
0,11
reaction.
The overall rate of hydrogenation (Figure 5)
1
decreased with increase in the liquid flow rate. In the range
of 10-25 mL/h liquid flow rate, the fall in the rate is fast,
and eventually it becomes slower. The model predictions for
the effect of liquid flow rate on the conversion and overall
(9) Telkar, M. M.; Rode, C. V.; Jagannthan, R.; Rane, V. H.; Chaudhari, R. V.
J. Mol. Catal. A: Chem. 2002, 187, 81.
280
•
Vol. 10, No. 2, 2006 / Organic Process Research & Development

Figure 7. Effect of hydrogen pressure on rate of hydrogena-
tion.Reaction conditions: liquid flow rate, 20 mL/h; H flow
2
rate, 20 L/h; temperature, 100 °C.
Figure 5. Effect of liquid flow rate on rate of hydrogenation.
Reaction conditions: pressure, 20 bar; H flow rate, 20 L/h;
temperature, 100 °C; catalyst weight, 20 g.
2
Figure 6. Effect of H
conditions: liquid flow rate, 20 mL/h; H
temperature, 100 °C.
2
pressure on conversion. Reaction
Figure 8. Effect of hydrogen pressure on the selectivity of B
and B D. Reaction conditions: liquid flow rate, 20 mL/h;
catalyst weight, 20 g; H flow rate, 20 L/h.
2
D
2
flow rate, 20 L/h;
1
2
B
1
D were exactly opposite to each other (see Figure 8). In
the lower range of H pressure (10-40 bar), the difference
between the selectivities of B D and B D was larger;
eventually, 50:50 formation of B D and B D could be
obtained at 70 bar H pressure. The predictions of the model
equations for the effect of pressure on the conversion of B D,
selectivity of B D and B D, and overall rate of hydrogenation
is in good agreement with the experimental results.
Effect of Hydrogen Flow Rate. The effect of hydrogen
gas flow rate on conversion of B D and selectivities of B
and B D were studied in the range of 20-80 L/h at constant
pressure, temperature, and liquid flow rate conditions. The
conversion of B D was slightly lower at the lower and higher
gas flow rates as compared with that in the middle range
Figure 9). The reason for this situation is that at lower gas
rate of hydrogenation are in well agreement with the
experimental results (see Figures 3 and 5).
2
2
1
Effect of H
on the conversion of B
0 bar H pressure for 10 and 20 g of catalyst loadings and
at 100 °C, and the results are shown in the Figure 6. Initially,
D conversion increased almost linearly with increase in
pressure up to 40 bar H pressure. The overall rate of
hydrogenation was also found to increase with increase in
pressure (Figure 7) for both the catalyst loadings. The
increase in H pressure increases the gas-liquid and liquid-
2
Pressure. The effect of hydrogen pressure
2
1
3
D was studied in the range of 10-
2
7
2
3
2
1
B
3
2
3
2
D
H
2
1
2
solid mass transfer coefficient, leading to higher conversion
and hydrogenation rates. It was interesting to note that with
increase in H pressure, the selectivity patterns of B D and
2 2
3
(
(
10) Malyala, V.; Rajashekharam, Jagannthan R.; Chaudhari, R. V. Chem. Eng.
Sci. 1998, 53, 787.
11) Chaudhari, R. V.; Jagannthan, R.; Mathew, S. P. AIChE J. 2002, 48, 110-
25.
flow rate the gas cannot overcome the gas-liquid mass
transfer resistance, while at higher gas flow rates, although
the gas reduces the liquid film thickness around the catalyst
(
1
Vol. 10, No. 2, 2006 / Organic Process Research & Development
•
281

Table 3. Kinetic parameters for the hydrogenation of B
3
D
rate constants adsorption constants
3
2
3
(
(m ) /kmol kg‚s)
(kmol/m )
temp (K)
k
1
k
2
K
B
K
C
3
3
3
3
3
33
43
53
63
73
1.940
3.178
5.037
7.802
11.804
1.210
2.790
6.140
12.934
26.178
8.101
9.80
11.728
13.898
16.32
2.11
2.60
3.183
3.843
4.594
with an increase in temperature, obviously because at higher
temperature B D gets converted to the saturated diol B D.
Reactor Modeling. A trickle-bed reactor model for the
hydrogenation of B D was developed using the rate equations
2
1
3
proposed by Telkar et al. based on the rate data in a slurry
reactor to represent the intrinsic kinetics for different reaction
9
steps as shown in Scheme 1.
Figure 9. Effect of gas flow rate on conversion of B
Reaction conditions: liquid flow rate, 20 mL/h; H pressure,
0 bar; temperature, 100 °C; catalyst weight, 10 g.
3
D.
2
wk A*B
1
l
2
r1 )
r2 )
(1)
(2)
2
(
1 + K B + K c )
B l c l
wk A*C
2
l
2
(
1 + K B + K c )
B
l
c
The overall rate of hydrogenation can be given as
wk A*(B + k C )
1
l
21 l
RA )
(3)
2
(
1 + K B + K c )
B l c
l l 2 2
where A, B , and, C represent the concentrations of H , B D,
and B D. A nonlinear least-squares regression analysis was
1
used to obtain the values of the kinetic parameters in the
above rate equation. For this purpose an optimization
1
2
program based on the Marquardt’s method was used. The
different kinetic parameter values evaluated for the above
rate equation are given in Table 3.
To develop a trickle-bed reactor model applicable to
hydrogenation of B
3
D, the approximate solution of the
catalytic effectiveness factor was performed for partial
13
wetting of catalyst particles. The overall catalyst effective-
ness factor could be expressed as a sum of the weighted
average of the effectiveness factor in the dynamic liquid-
covered, stagnant liquid-covered and complete gas-covered
zones and with the assumptions that (a) gas and liquid phases
are in plug flow; (b) the liquid-phase reactant is nonvolatile
and was in excess as compared to the gaseous reactant; (c)
the gas-liquid, liquid-solid, and intraparticle mass transfer
Figure 10. Effect of temperature on conversion of B
selectivity of B D and B D. Reaction conditions: liquid flow
flow rate, 20 L/h; pressure, 20 bar; catalyst
3
D and
1
2
rate, 20 mL/h; H
weight, 10 g.
2
pellets, the gas phase becomes continuous, and the liquid
phase becomes dispersed. With further increase in gas flow
rate, the flow regime changes to spray flow regime, and this
causes greater velocities in the liquid phase by increasing
the drag forces on the liquid phase. The selectivities to both
B D and B D were nearly constant since the change in gas
2 1
flow rate did not affect the intrinsic reaction kinetics to a
considerable extent.
resistances for H are considered, whereas the liquid-solid
2
and intraparticle mass transfer resistances for the liquid phase
were negligible; (d) the interphase and intraparticle heat
transfer resistances were negligible. The catalyst effectiveness
factor equation for the hydrogenation of B
3
D could be
Effect of Temperature. The effect of temperature on
developed on the basis of the approaches already reported
conversion of B
were studied by varying the temperature from 60 to 100 °C
Figure 10). The conversion increased from 5 to 30% with
3 2 1
D and selectivity patterns of B D and B D
1
4,15
in the literature.
Under the conditions of significant
(
(12) Marquardt, D. W. J. Sco. Ind. Appl. Math. 1963, 11, 431.
increase in temperature from 60 to 100 °C since, the temper-
ature has a greater impact on the reaction kinetics. The
(13) Tan, C. S.; Smith, J. M. Chem. Eng. Sci. 1980, 35, 1601
(14) Ramachandran, P. A.; Chaudhari, R. V. Three Phase Catalytic Reactors;
Gordan and Breach: New York, 1983.
1 2
selectivity patterns of B D and B D showed opposite trends
(15) Bischoff, K. B. AIChE J. 1965, 11, 351.
282
•
Vol. 10, No. 2, 2006 / Organic Process Research & Development

intraparticle gradient for the gas-phase reactant (H
when the liquid-phase reactant was in excess, the overall
rate of hydrogenation of B D was given as
2
) and
B
2
D in liquid phase in terms of dimensionless parameters.
dpld
C r 21 l
η R k c ø
3
)
(10)
2
dz
q (1 + b k + c k )
B
l
b
l c
η wk A*(B + k C )
C
1
l
21
l
RA )
(4)
2
(
1 + K B + K c )
This equation represents the change in the concentration of
D in liquid phase in terms of dimensionless parameters,
B
l
c
B
1
where η
C
, the overall effectiveness factor for the spherical
where
catalyst particle was given as
f a
f a
s ld
d
ld
ø )
+
+
1
φ
1
3φ
2
2
2
{
ηC ) coth 3φ -
(5)
(1 + η φ /Ν ) (1 + η φ /N + η φ /R N )
C δ C s C s s
(
)
(
1 - f - f ) a
d s ld
φ is the Thiele parameter, and final dimensionless parameters
are given in Appendix 1. For calculating φ, the values of
(11)
2
}
(
1 + η φ /Ν )
C g
effective diffusivity were estimated using the standard
A programe code in Q-Basic was developed to get the
output of eqs 7-10. These equations were solved using a
fourth-order Runge-Kutta method with the following initial
conditions:
correlation.1
5,16
Various correlations used to calculate the
different coefficients in this work are given in Appendix 1.
The saturation solubility of hydrogen was calculated as
(
A*) ) [P - (P ) ](H )
(6)
T
V T
e T
At z ) 0, a ) b ) 1; c ) P ) 0
(12)
l
l
l
l
v
where, P is the vapor pressure of solvent.
The model equations developed above, allowed the prediction
of concentrations of products/reactants along the length of
the reactor. At any given length of the reactor the fractional
At steady-state conditions, the sum of the convection term
and the gas-liquid mass transfer term were in equilibrium
with the liquid-solid mass transfer term in the dynamic zone
and the volumetric mass exchange between the dynamic and
stagnant zone. The final mass balance equations in dimen-
conversion of B
3
D (X
B
) could be given as
X
) 1 - b
(13)
(14)
B
l
sionless form for species A (H
2
) are given as
The overall rate of hydrogenation was calculated as
dald
dz
η R (b + k c )
C
r
l
21 l
)
R (1 - a ) -
x
2
Ul
gl
ld
(
1 + b k + c k )
RA ) (C + 2 P )
l
b
l
c
l
l
L
f a
d
ld
+
l
Here U , is the liquid velocity in m/s, L is the length of the
2
(
1 + η φ /N )
C
d
catalyst bed in m, C
l
, P
l
are the concentrations of B
2
D and
(
7)
f a
s
ld
B
1
D, respectively. The selectivities of B
2
D and B
1
D were
{
2
2
}
(
1 + η φ /N + η φ /R N )
calculated as
C
s
c
s s
cl
Similarly, the mass balance of liquid-phase reactant/products
in dimensionless form can be given as
S
S
)
)
× 100
× 100
(15)
(16)
B D
2
(1 - bl)
Pl
dbld
η R b ø
C
r l
B D
)
-
(8)
1
(1 - bl)
2
dz
q (1 + b k + c k )
B
l
b
l c
The applicability of the model was verified by comparing
the predicted concentration vs time profiles as well as
conversion of B D and selectivities to B D and B D with
3 2 1
This equation represents the change in the concentration of
B
3
D in liquid phase in terms of dimensionless parameters.
the experimental results under various inlet conditions. These
results have already been discussed above and are shown in
Figures 5-8, which showed an excellent agreement between
the predicted values and the experimental data.
dcld
η R (b - k c ) ø
C
r
l
21 l
)
(9)
2
dz
q (1 + b k + c k )
B
l
b
l c
This equation represents the change in the concentration of
Conclusions
(
(
(
(
(
(
(
16) Wilke, C. R.; Chang, P. AIChE J. 1955, 1, 246.
17) Satterfield, C. N.; Way, P. F. AIChE J. 1972, 18, 305.
18) Goto, S.; Smith, J. M. AIChE J. 1975, 21, 706.
19) Satterfield, C. N.; Vab Eek, M. W.; Bliss, G. S. AIChE J. 1978, 24, 709.
20) Zheng, L. P.; Smith, J. M.; Herskowitz, M. AIChE J. 1984, 30, 500.
21) Hochmann, J. M.; Effron, E. Ind. Eng. Chem. Fund. 1969, 8, 63.
22) Sato, Y.; Hiros, T.; Takahashi, F.; Toda, M. J. Chem.. Eng. Japan 1973, 6,
Hydrogenation of 2-butyne-1,4-diol (B
using 1% Pt/CaCO catalyst in a fixed-bed reactor. This
process represents a consecutive reaction scheme giving
-butene-1,4- and butane-1,4-diols. In a continuous hydro-
3
D) was carried out
3
2
genation process, the ratio of butene- and butane diols could
be manipulated by tailoring the operating conditions for the
same catalyst. The liquid flow rate and pressure of the
hydrogen gas showed a significant effect on the conversion
1
47.
(
(
23) Zai-Sha, M.; Tian-Ying, X.; Chen, J. Chem. Eng. Sci. 1993, 48, 2697.
24) Stephen, H.; Stephen, J. Solubilities of Inorganic and Organic Compounds;
Pergamon Press: U.K. 1963; Vol. I, Part I.
(
25) Mills, P. L.; Dudukovic, M. P. AIChE J. 1981, 27, 893.
Vol. 10, No. 2, 2006 / Organic Process Research & Development
•
283

3
of B D and overall rate of hydrogenation. A theoretical model
N
g
Nusslet number for the gas phase
was also developed incorporating the conditions of external
and intraparticle mass transfer, partial wetting of the catalyst
and reaction kinetics of butyne diol hydrogenation in a batch
slurry reactor. The reactor model was validated by carrying
out hydrogenation experiments under various reactor inlet
conditions, and model predictions were found to agree well
with the observed conversion, selectivity data.
qB
stoichiometric ratio () B /A*)
li
r
1
, r
2
reaction rate for the individual hydrogenation steps,
3
kmol/m /s
R
R
radius of the pellet, m
3
A
overall rate of hydrogenation, kmol/m /s
Re
l
Reynolds number for the liquid phase
external surface area of the catalyst pellet, m2
gas velocity (superficial), m/s
liquid velocity (superficial), m/s
weight of catalyst, kg/m3
S
ex
U
g
l
NOTATION
U
a
a
a
a
a
l
l
concentration of hydrogen in liquid phase, () A /A*),
dimensionless
W
Z
reactor length, m
s
concentration of hydrogen on the catalyst surface,
) A /A*), dimensionless
B P
external surface area of the pellet [) 6(1 - ꢀ )/d ],
(
s
Greek letters
p
t
R
gl
R
ls
R
r
ꢀ
dimensionless gas-liquid mass transfer coefficient
dimensionless liquid-solid mass transfer coefficient
dimensionless reaction rate constant
porosity of catalyst
-
1
m
packing external surface area per unit volume of the
reactor [) Sex(1 - ꢀ )/V
b
B
]
catalyst area wetted, m-1
w
ꢀ
b
bed voidage
A
A
l
concentration of hydrogen in liquid phase, kmol/m3
s
concentration of hydrogen on the catalyst surface,
kmol/m3
Appendix 1
(1) Dimensionless parameters used in the model:
A*
concentration of hydrogen in equilibrium with liquid,
kmol/m3
b
l
concentration of B
dimensionless
3
l
D in liquid phase () B /Bli),
D in liquid phase, kmol/m3
D in liquid phase, kmol/m3
D in liquid phase, () C
gas-liquid mass transfer
liquid-solid mass transfer
gas-solid mass transfer
Nusslet no. in dynamic zone
Nusslet no. in stagnant zone
reaction rate constant
a
a
a
N
N
g
) k
l
a
B
L/U
L/U
L/U
) R ksd/3 D
) R kSs/3 D
) Lwk /U
l
B
l
concentration of B
initial concentration of B
3
ls ) k
S
a
p
a
l
B
li
3
gs ) kgs
p
l
c
l
concentration of B
dimensionless
2
l
/Bli),
d
e
S
e
R
r
l
b
i
l
C
l
concentration of B
2
D in liquid phase, kmol/m3
equilibrium constants
k
) k /k ; k ) K B ;
2
1
k
2
1
b
B
l
C
) K
C
B
l
2
D
e
effective diffusivity, m /s
1
/2
Thiele parameter
S (k B + k C )
R
3
p 1 l 2 l
2
D
M
molecular diffusivity, m /s
φ )
2
[
]
D ( 1 + K B + K C )
e
B
l
C l
d
p
particle diameter, m
f
f
f
d
fraction of catalyst wetted by dynamic liquid
s
fraction of catalyst wetted by the stagnant liquid
wetted fraction
(2) Correlations used in this work:
w
F
k
liquid flow rate, kg/h
3
3
parameter
author(s)
ref
1
, k
2
c
reaction rate constants, (m /kg) (m /kmol‚s)
dimensionless rate constant () k /k
B c
dimensionless equilibrium constant (k ) K Bli; k )
k
21
2
1
)
molecular diffusitivity
gas-liquid mass
Wilke and Chang
Goto and Smith
16
18
k
b
s
, k
b
transfer coefficient
liquid-solid mass
transfer coefficient
gas-particle mass transfer
coefficient (value used)
volumetric mass
C li
K B )
Satterfield et al.(1978)
Zheng Lu et al.
19
20
21
liquid-solid mass transfer coefficient, m s-
gas-particle mass transfer coefficient, m s-1
1
k
k
gs
3
K
A
, K
B
, K
C
adsorption constants, m /kmol
Hochmann and Effron
gas-liquid mass transfer coefficient, s-
1
exchange coefficient
total liquid hold-up
static liquid hold-up
(value used)
saturation solubility
wetting efficiency
k
l
a
B
Sato et al.
22
23
K
ex
exchange coefficient between dynamic and stagnant
Zai-sha Mao et al.
liquid, s-1
Stephen and Stephen
Mills and Dudukovic
24
25
L
reactor length, m
N
d
s
Nusslet number for the liquid phase in the dynamic
zone
Received for review October 28, 2005.
OP050216R
N
Nusslet number for the liquid phase in the stagnant
zone
284
•
Vol. 10, No. 2, 2006 / Organic Process Research & Development