4424
Ind. Eng. Chem. Res. 1996, 35, 4424-4433
Kinetics of the Liquid Phase Hydrogenation of Di- and Trisubstituted Alkylbenzenes over a Nickel Catalyst S. Toppinen,† T.-K. Rantakyla1 ,‡ T. Salmi,*,§ and J. Aittamaa† Laboratory of Industrial Chemistry, A° bo Akademi, FIN-20500 A° bo, Finland, Department of Process Engineering, University of Oulu, FIN-90570 Oulu, Finland, and Neste Engineering, P.O. Box 310, FIN-06101 Porvoo, Finland
The liquid phase hydrogenation kinetics of five di- and trisubstituted alkylbenzenes, xylenes, mesitylene, and p-cymene were determined in a semibatch reactor operating at hydrogen pressures of 20-40 atm and at temperatures of 95-125 °C. Commercial preactivated catalyst particles of nickel-alumina were used in all experiments. The hydrogenation activity of the aromatic compound was affected by both the number of substituents and their relative positions in the benzene ring. The trisubstituted benzene (mesitylene) had a lower reaction rate than the disubstituted compounds (xylenes). The activity of the different substituent positions decreased in the order para > meta > ortho. The main reaction product was always the completely hydrogenated cycloalkane; no cycloalkenes were detected. The hydrogenation rates were virtually constant at low and intermediate conversions of the aromatics, but at high conversions the rates decreased. Rate equations based on the formation of partially hydrogenated surface complexes were derived, and the kinetic parameters were estimated from a heterogeneous reactor model with nonlinear regression analysis. The rate equations were able to describe the features of the experimental data. Introduction The present work concerns the catalytic liquid phase hydrogenation of di- and trisubstituted aromatic molecules to their cyclic homologues. The overall reaction can be written as follows:
where R1, R2, and R3 denote alkyl groups or hydrogen. This study is a continuation of the work initiated with the hydrogenation of benzene and some monosubstituted alkylbenzenes (Toppinen et al., 1996). The industrial background of the actual research lies in the production of speciality chemicals, in the removal of aromatics from solvents, and in the production of aromatic-free fuels. Even though the principal products of the hydrogenation reaction are well knownsthe completely saturated cyclic compound is typically the dominating main product and some cycloalkenes are obtained as byproductss the kinetics and mechanism of the liquid phase hydrogenation of substituted aromatics have not been thoroughly investigated. Numerous studies have been reported for gas phase hydrogenation of benzene and monosubstituted benzenes, as reviewed, e.g., by Lindfors et al. (1993). However, most industrial dearomatization processes are carried out in the liquid phase at high pressures (e.g., at 20-40 atm) by using porous catalyst pellets. Thus, the gas-liquid mass transfer and the mass transfer in the porous catalyst pellets put their * Address correspondence to Dr. T. Salmi, Laboratory of Industrial Chemistry, A° bo Akademi, Biskopsgatan 8, FIN20500 A° bo, Finland. Tel.: +358-2-2654427. Fax: +358-22654479. E-mail:
[email protected]. † Neste Engineering. ‡ University of Oulu. § A ° bo Akademi.
S0888-5885(95)00636-1 CCC: $12.00
fingerprints on the experimentally observed data. This fact makes the interpretation of the kinetic data and the subsequent development of rate equations difficult and ambiguous. In spite of these obstacles, there exists always a need for rate equations, which can be used in reactor design. The present study aims to demonstrate the experimental determination of the precise kinetics and estimation of rate parameters for the liquid phase hydrogenation of some di- and trisubstituted aromatics, such as o- m-, and p-xylenes (1,2-, 1,3-, and 1,4-dimethylbenzenes), p-cymene (1-isopropyl-4-methylbenzene), and mesitylene (1,3,5-trimethylbenzene), over a supported nickel catalyst. Experimental Section The kinetic experiments were carried out in a computer-controlled reactor system manufactured by Xytel B.V. Europe. The reactor was operated by an integrated control system, which was run under a PC-DOS microcomputer. The PC computerized all the main control and monitoring actions, providing also data acquisition and scheduling facilities. The reactor (inner diameter 100 mm, total volume 1000 mL) was operated in a semibatch mode, i.e., hydrogen was fed into the reactor to maintain the pressure constant during the experiments. The desired total pressure was obtained by adjusting the pressure of the hydrogen supply. Hydrogen flow rates were regulated with a mass flowmeter (Bronkhorst High-Tech F 231C-FA). The reactor was equipped with a magnetic propeller stirrer, whose speed motor was installed with computer monitoring and setpoint adjusting. The reactor was heated in a three-zone electric heating furnace, and the temperatures of the heating jackets were controlled by a computer. The reactor was equipped with two internal thermocouples. The catalyst was placed in a static basket. The three liquid phase hydrogenation of aromatics was studied on a commercial nickel-alumina catalyst (Crosfield; Ni 16.6 wt %, specific surface area 108 m2/g, © 1996 American Chemical Society
Ind. Eng. Chem. Res., Vol. 35, No. 12, 1996 4425 Table 1. Selectivity of the Hydrogenation Reaction for the cis-Isomer (%) at Different Reactor Temperatures and Pressures 20 bar 1,2-dimethylcyclohexane 1,3-dimethylcyclohexane 1,4-dimethylcyclohexane 1,3,5-trimethylcyclohexane 1-isopropyl-4-methylcyclohexane
95 °C
125 °C
54.3 77.3 47.0 80.4 54.9
53.3 77.2 44.6 78.3 54.3
mean pore volume 0.37 cm3/g, bulk density 0.86 g/cm3). Hydrogen was supplied by AGA (>99.9995%). The model compounds used in these experiments were o-xylene (>99%, Merck-Schuchardt), m-xylene (>99%, J.T. Baker), p-xylene (>99%, Merck-Schuchardt), pcymene (>95%, Fluka), and mesitylene (>99%, MerckSchuchardt). Samples were analyzed with a gas chromatograph (Hewlett-Packard 5830A) equipped with a 30 m long capillary column (J&W Scientific, DB-624) and a flameionization detector (FID). The column temperature was 100 °C. The identification of hydrogenated products was done by gas chromatography-mass spectrometry (GCMS) analysis. Kinetic experiments were carried out at pressures of 20 and 40 bar and at temperatures of 95 and 125 °C. In addition to this, the hydrogenation of mesitylene was also investigated at 10 and 30 bar and at 140 °C. The amount of catalyst was 10 g in the experiments performed at 125 °C, and it was increased to 20 or 30 g when operating at the lowest temperature. The particle size of trilobe catalyst extrudates used in these experiments was as follows: the mean particle length was 3 mm, and the radius of the lobe was 0.25 mm. Activation of the nickel catalyst was performed before each experiment in order to maintain the catalyst activity at a constant level. The activation was done at atmospheric pressure in flowing hydrogen (3 dm3/ min). During the activation, the catalyst was heated to 120 °C and maintained at this temperature for 1 h, after which the temperature was increased to 260 °C at a rate of 80 °C/h. The hydrogen flow was continued at 260 °C for 1.5 h. The reactor was then cooled down to 80 °C under a small hydrogen flow (0.2 dm3/min). The activation procedure was carried out automatically using the recipe and scheduling programs available in the control system package. Therefore, the possible differences between activation treatments were eliminated. Typically, the pure aromatic reagent (335 mL) was loaded into the reactor through the feeding tube by using hydrostatic pressure of the liquid. Based on preliminary tests, a stirrer speed of 1900 rpm was chosen to be suitable for the elimination of external mass transfer resistances. Temperature and pressure values were adjusted to desired values, and after 20 min the first sample was taken for GC analysis. Sampling intervals were approximately 30 min, but when the aromatic concentration was lower than 15%, more frequent intervals were used (10-15 min). The duration of experiments varied from 3 to 7 h, depending on conditions and amounts of the catalyst. The temperature was controlled within (4 °C of the desired value and the pressure varied within (0.3 bar, but at the end of experiment, when the aromatic concentration was low, the pressure increased about 10% from the desired value. After the experiment, the reactor was drained and rinsed using nitrogen flow and heat treatment.
40 bar 140 °C
76.7
90 °C
125 °C
55.2 76.4 48.4 79.8 52.2
54.4 76.3 47.5 78.1 52.4
140 °C
77.9
Figure 1. Production of cis- and trans-isomers during the hydrogenation experiments at 125 °C and 40 bar.
Results and Discussion Qualitative Considerations. The identification of hydrogenated products was done by GC-MS. The results revealed that all the di- and trisubstituted alkylbenzenes had two main hydrogenated products, the cis- and trans-isomers of the corresponding cyclohexanes. The hydrogenated products of o-, m-, and pxylene, p-cymene, and mesitylene were 1,2-dimethylcyclohexanes (cis- and trans-isomers), 1,3-dimethylcyclohexanes, 1,4-dimethylcyclohexanes, 1-isopropyl-4methylcyclohexanes, and 1,3,5-trimethylcyclohexanes, respectively. The relative selectivities of the formation of cyclohexanes in cis- and trans-configuration are shown in Table 1. Alkylbenzenes with substituents in para- and orthopositions were hydrogenated to different stereoisomers with almost equal rates in all conditions studied. The cis-isomers were found to be slightly more dominant product in the experiments with o-xylene and p-cymene. To the contrary, slightly higher amounts of hydrogenated products of p-xylene existed in trans-form. Alkylbenzenes with substituents in meta-positions reacted dominantly to corresponding cyclohexanes in cis-position. The selectivity for cis-formation was found to be approximately 80% in all conditions studied. According to literature data (Rylander, 1979), in addition to the choice of the catalyst metal, the temperature was the most important parameter influencing the stereochemistry of aromatic hydrogenation. All three xylenes appeared to yield on hydrogenation at room temperature products having predominantly the cis-configuration, but in the case of o- and p-xylene, the proportion of the trans-isomers increases with temperature, and above 150 °C these may be the major isomers. Since the cis-form of 1,3-dimethylcyclohexane is the more stable, the yield of the trans-isomer does not
4426 Ind. Eng. Chem. Res., Vol. 35, No. 12, 1996
a
d
b
e
c
Figure 2. Model I: measured (symbols) and estimated (curves) aromatic concentrations of o-xylene (a), m-xylene (b), p-xylene (c), mesitylene (d), and p-cymene (e) experiments.
increase with temperature (Bond, 1962). Table 1 shows that our experimentally observed results are consistent with those presented in the literature. The ratio between the cis- and trans-forms of the product remained constant during the experiments (Figure 1), i.e., the selectivity was independent of the aromatic concentration. The measured concentration trends in the hydrogenation experiments are illustrated in Figure 2 using t′ ) tmcat/VL as the abscissa. In that way, the experiments
with different amounts of catalyst and different liquid loadings become directly comparable. If the evaporation of the liquid into the gas phase is neglected, the mass balance of a liquid phase component is
dnLi ) mcatri dt
(2)
where the generation rate (ri) is given per catalyst mass (mcat). Furthermore, if the volume of the liquid phase
Ind. Eng. Chem. Res., Vol. 35, No. 12, 1996 4427 Table 2. Average Reaction Rate at Different Temperatures and Pressures [mol/(h kg)] 20 bar 1,2-dimethylbenzene 1,3-dimethylbenzene 1,4-dimethylbenzene 1,3,5-trimethylbenzene 1-isopropyl-4-methylbenzene
95 °C
125 °C
16.5 17.8 25.7 20.7 19.2
45.6 52.7 60.2 37.7 40.4
67.6
90 °C
125 °C
29.2 29.0 54.5 24.3 35.3
77.1 91.3 106.6 65.3 60.2
140 °C
96.2
as follows (model I):
(VL) is constant, eq 2 can be simplified to
dcLi ) ri dt′
40 bar 140 °C
(3)
i.e., the reaction rates are proportional to the slopes of the curves in Figure 2. The catalyst amount was 10 g in all the experiments at 125 °C and 20 g (with mesitylene and p-cymene) or 30 g (with xylenes) in the experiments at 95 °C. The liquid volume varied between 320 and 335 mL. To compare the effects of the temperature and pressure on the reaction rate, average reaction rates were calculated using the concentration interval 10-80 wt % (Table 2). At 95 °C, the differences between the observed reaction rates of the different aromatics were so small that they could not be distinguished from the experimental error properly. At 125 °C, mesitylene seems to have a lower reaction rate than the xylenes (Table 2). This indicates that the reaction rate decreases with an increase in the number of substituents in the benzene ring. In our previous work (Toppinen et al., 1996), it was shown that the reaction rate decreased with the increasing length of the substituent. A comparison of the reaction rates of xylenes (Table 2) reveals that the relative positions of the substituents also have a significant effect on the reaction rate. The para-position seems to have the highest and the ortho-position the lowest reactivity. The reaction rates of mesitylene and p-cymene were of the same order of magnitude. This can be the result of opposite effects of the bulky isopropyl substituent in the p-cymene molecule and, on the other hand, the greater number of substituents and their meta-positions in the mesitylene molecule. The apparent reaction order with respect to the liquid phase hydrogen concentration, which is proportional to the pressure, seems to be little less than one. The apparent activation energy varies between 25 and 45 kJ/mol. Modeling of the Hydrogenation Kinetics. Although the present data (experiments were made only at two temperatures and pressures) are probably insufficient for a profound comparison of reaction mechanisms, two kinds of reaction mechanisms, leading to different types of rate expressions, were considered in order to obtain rate equations for practical use: first, a mechanism where hydrogen and the aromatic compound are adsorbed competitively on the surface of the catalyst and hydrogen is added to the aromatic ring in three sequential surface reaction steps (Smeds et al., 1995), and second, a mechanism where the reaction proceeds via an intermediate surface complex (Temkin et al., 1989). Smeds et al. (1995) studied the gas phase hydrogenation of ethylbenzene and proposed a mechanism where aromatic and hydrogen molecules are competitively adsorbed on the vacant catalyst active sites. Hydrogen is added to the aromatic molecule in three sequential surface reaction steps. The mechanism can be described
k1
z AH2* + 2* A* + 2H*y\ k
(4)
-1
k2
yk z AH4* + 2* AH2* + 2H* \
(5)
-2
K
AH4* y\z AH′4* k3
AH4* + 2H* 98 AH6* + 2* k′3
AH′4* + 2H* 98 AH′6* + 2* KA
A+*\ y z A*
(6)
(7)
(8)
(9)
KH
y z 2H* H2 + 2* \ rapid
AH6* 98 AH6 + * rapid
AH′6* 98 AH′6 + *
(10)
(11) (12)
where the hydrogen addition steps are assumed to be rate determining. The surface intermediates AH2*, AH4*, and AH′4* are supposed not to be the substituted cyclohexadiene and cyclohexane compounds but instead compounds which preserve their aromatic character (Smeds et al., 1995). Particularly cyclohexadiene is an improbable intermediate because the hydrogenation rate of this compound is known to be very high. The surface intermediate AH4* is isomerized to AH′4*, which reacts forward to produce the trans-form of the product, AH′6. In the derivation of the rate equation, we assume that the surface reactions are rate determining, whereas the adsorption steps of hydrogen and the aromatic compound are rapid enough for the quasi-equilibrium hypothesis to be applied. The stationary state assumption for the adsorbed surface intermediates AH2*, AH4*, and AH′4* gives the expressions for the production rates of the cis- and trans-configurations:
R)
R′ )
k1k2k3KAKH3cAcH23 R[(β/R)KAcA + xKHcH2 + 1]3 k1k2Kk′3KAKH3cAcH23 R[(β/R)KAcA + xKHcH2 + 1]3
(13)
(14)
4428 Ind. Eng. Chem. Res., Vol. 35, No. 12, 1996
where
steps (model II):
R ) k2k′KH2cH2 + k′k-1KHcH + k-1k-2
k1
A* + H2 y\ z A(H2)* k
(15)
(23)
-1
k2
β ) [(K + 1)k1k2 + k1k′ + k2k′]KH2cH2 + (k1k-2 + k′k-1)KHcH + k-1k-2 (16)
z A(H2)2* A(H2)* + H2 y\ k
(24)
-2
k3
k′ ) k3 + Kk3′
A(H2)2* 98 AH4*
(17)
(25)
rapid
AH4* + H2 98 AH6*
The ratio between the production rates of the stereoisomers is simply
R/R′ ) k3/Kk3′
(18)
The experimental results showed that the hydrogenation reactions are practically irreversible in actual conditions, thus the backward rate constants were set to zero in order to avoid the overparametrization of the model (k-1 ) k-2 ) 0). In addition, all rate constants were assumed to be equal (k1 ) k2 ) k3 ) k′3):
R)
R′ )
k1KAKHcAcH (3KAcA + xKHcH + 1)3(K + 1) k1KKAKHcAcH (3KAcA + xKHcH + 1)3(K + 1)
(19)
(20)
As well, only the first hydrogenation step (4) could be assumed to be rate determining, in which case the term 3KA should be replaced with KA in the denominators of eqs 19 and 20. The stereoselectivity was observed to be independent of the concentrations and also, at the studied temperature and pressure interval, rather independent of the reaction conditions. Therefore, the kinetic parameters could be estimated using the expression for the total hydrogenation rate:
Rtot ) R + R′ )
k1KAKHcAcH (3KAcA + xKHcH + 1)3
(21)
The temperature dependence of the adsorption coefficients KA and KH was presumed to be negligible, and the rate constant k1 was assumed to depend on the temperature according to the Arrhenius’ law:
[ (
k1 ) k1Tref exp -
)]
Ea 1 1 R T Tref
(22)
where a temperature of 100 °C was used as the reference temperature Tref. Temkin et al. (1989) have proposed a mechanism where the aromatic molecule forms an intermediate complex with hydrogen. The complex is then isomerized to adsorbed cyclohexene and rapidly hydrogenated to cyclohexane. The aromatic molecules are adsorbed on the catalyst surface by replacing adsorbed cyclohexane molecules. If the stereochemistry is neglected, the mechanism can be described with following reaction
(26)
K4
A + AH6* y\z A* + AH6
(27)
After applying the stationary state assumption for the surface components and assuming the quasi-equilibrium of the desorption step (eq 27), the rate of the overall reaction becomes
[
R ) k3 1 +
k-2 + k3 k-1k-2 + k-1k3 + k2k3cH × + k2cH kkc 2
(
1 2 H
1+
cAH6 K4cA
)]
-1
(28)
The rate expression of model II was simplified by assuming the first two reaction steps (eqs 23 and 24) to be irreversible (k-1 ) k-2 ) 0) and the second reaction to be rapid compared to the first reaction (k1 , k2). The simplified rate expression is thus
R)
k1cAcH (k1/k3)cAcH + (1/K4)cAH6 + cA
(29)
The temperature dependence of the adsorption coefficient K4 was assumed to be negligible, and the rate constants k1 and k3 were assumed to depend on the temperature according to the Arrhenius’ law and have equal activation energies. Reactor Model. In our previous study (Toppinen et al., 1996), which concerned the hydrogenation of monosubstituted alkylbenzenes, it was shown by rigorous simulations that the mass transfer resistance inside the porous catalyst particles influenced the overall hydrogenation rate. Because the behaviors of the di- and trisubstituted alkylbenzenes are similar to those of the monosubstituted benzenes, the heterogeneous reactor model was used in the present work. The fundamental assumptions concerning the model were as follows: (i) both the bulk phases and the catalyst particles are in dynamic conditions; (ii) the bulk phases and the catalyst particles are principally nonisothermal; (iii) the simply Fickian law can be described for the fluxes in the gas and liquid films as well as in the catalyst particle; and (iv) the hydrogen feed into the reactor can be described with a proportional (P) controller. The mass balances for the gas and liquid bulk phases can be written as follows:
dnGi ) -VRNGLia + Fi dt
(30)
dnLi ) VRNGLia - VRNLSiaS dt
(31)
where NGLi and NLSi denote the fluxes from the gas bulk
Ind. Eng. Chem. Res., Vol. 35, No. 12, 1996 4429
to the liquid bulk and from the liquid bulk into the catalyst particles, respectively. The bulk phase mass balances are coupled to those of the catalyst particles. For the cylindrical geometry, the component mass balance has the following form in transient conditions:
(
)
∂ci Di,eff ∂2ci 1 ∂ci F ) + + ri 2 2 ∂t x ∂x Rp ∂x
(32)
Because the bulk phase temperature was recorded experimentally, there was no need to use the energy balance. Instead, the time course of the temperature was obtained from a spline function fitted to the experimental temperature. Inside the catalyst particle, the temperature could not be recorded experimentally. Therefore, the dynamic energy balance was used in the form
(
)
(-∆HR) λ ∂2T 1 ∂T ∂T ) + + Rtot 2 2 ∂t x ∂x cp FcpRp ∂x
(33)
The mass balances for the liquid bulk and the catalyst particle are coupled through the flux NLSi at the outer surface of the catalyst. The flux is obtained from the concentration gradient
( )
Di,eff ∂ci Rp ∂x
NLSi )
x)1
(34)
The fundamental geometrical relationships are used to express the term VRaS in eq 31 with the mass of the catalyst in the reactor (mcat) and density of the catalyst particle (F). It can be easily shown that
2mcat VRaS ) FRp
(36)
T|x)1 ) TL
(37)
In the center of the particle, concentration and temperature gradients disappear:
)0
(38)
∂T )0 ∂x x)0
(39)
x)0
The effective diffusion coefficients were computed from the molecular diffusion coefficients using the experimentally determined value for the porosity ( ) 0.5). The tortuosity factor (τ) was approximated as 4.0, which is the typical value for alumina (Satterfield, 1970).
Di,eff ) Di τ
(xj/D°ij) ∑ j*i
(41)
where the diffusion coefficients at infinite dilution (D°ij) were calculated using the Wilke-Chang equation (Reid et al., 1987). The catalyst density (F) was 1300 kg/m3, and the thermal conductivity (λ) was 0.15 W m-1 K-1. The effective particle radius (Rp) was 0.25 mm. The molar enthalpy of the liquid was calculated using standard molar enthalpies of formation and ideal heat capacities of the components and corrected with the Soave-Redlich-Kwong equation of state (Graboski and Daubert, 1978). The reaction enthalpy (-∆HR) was then obtained from the partial molar enthalpies. Typically the values of -∆HR ranged from 205 to 215 kJ/ mol. An average value of 210 kJ/mol was used as a constant in the reactor simulations. The hydrogen feed to the reactor was obtained from
FH2 ) Kp(VR - VL - VG)
(42)
Fi ) 0, i * H2 where the volumes of the liquid and gas phase (VL and VG, respectively) were calculated using the experimentally recorded reactor pressure and temperature. The molar volume of the liquid was obtained from the Chueh-Prausnitz equation (Reid et al., 1987) and the molar volume of the gas from the Soave-RedlichKwong equation of state. The flux between the liquid and gas bulk phases was obtained from the film theory (i.e., gas and liquid film coupled in series):
(cGi/Ki′) - cLi (1/kGiaKi′) + (1/kLia)
(43)
(35)
ci|x)1 ) cLi
| |
Di ) (1 - xi)/
NGLia )
is valid for cylindrical geometry. Since the external mass and heat transfer resistances around the catalyst particles were assumed to be negligible, the boundary conditions at the surface of the particle are
∂ci ∂x
The molecular diffusion coefficients were obtained from the equation (Wilke, 1950; Reid et al., 1987)
(40)
where kLia and kGia denote the volumetric mass transfer coefficients of the liquid and the gas film, respectively, and the vapor-liquid equilibrium constants are
Ki′ ) Ki
cGtot cLtot
(44)
The equilibrium constants (Ki) were estimated from the Soave-Redlich-Kwong equation of state (Graboski and Daubert, 1978). The complete mathematical model thus consisted of the ordinary differential equations (ODEs) (30 and 31) coupled to the parabolic partial differential equations (PDEs) (32 and 33). The coupled ODE-PDE system was transformed to ODEs, to an initial value problem, by discretizing the catalyst particle coordinate (x) with central difference formulas. Typically, a 5-point central difference formula was used. The ODE system created was the solved with the backward difference method (Gear’s method) implemented in the software LSODE (Hindmarsh, 1983). Parameter Estimation. In the estimation of the kinetic parameters, the mathematical model was coupled to a parameter estimation routine, which was used to minimize the difference between the observed and predicted concentrations (in wt %). Thus, the objective function was
RSS )
∑(wA,est - wA,obs)2
(45)
4430 Ind. Eng. Chem. Res., Vol. 35, No. 12, 1996
a
d
b
e
c
Figure 3. Model II: measured (symbols) and estimated (curves) aromatic concentrations of o-xylene (a), m-xylene (b), p-xylene (c), mesitylene (d), and p-cymene (e) experiments. Table 3. Estimated Parameters of Model I (Eq 21) parameter Rref, mol/(s kg) Ea, kJ/mol KA × 105, m3/mol KH × 104, m3/mol RRMSc RSS total RSS a
o-xylene 0.001a
0.020 ( 45.4 ( 1.6 9.0 ( 2.2 29.9 ( 22.8 1.46 69 1242
m-xylene
p-xylene
mesitylene
p-cymene
0.023 ( 0.001 52.1 ( 1.8 9.5 ( 2.3 18.8 ( 17.0 1.52 62
0.043 ( 0.006 35.8 ( 3.2 7.1 ( 2.9 2.4b 1.77 79
0.016 ( 0.002 29.8 ( 4.0 14.6 ( 8.7 129.7b 4.37 669
0.026 ( 0.004 21.3 ( 4.1 8.2 ( 3.3 1.7b 3.48 363
95% confidence interval. b Very large confidence interval. c Residual root mean square.
Ind. Eng. Chem. Res., Vol. 35, No. 12, 1996 4431 Table 4. Estimated Parameters of Model II (Eq 29)
a
parameter
o-xylene
m-xylene
p-xylene
mesitylene
p-cymene
k1(Tref), mol/(s kg) Ea, kJ/mol k3(Tref), mol/(s kg) K4 RRMS RSS total RSS
2.2 ( 0.2 46.0 ( 1.2 0.05 ( 0.01 14.5 ( 6.2 1.11 39 1089
2.2 ( 0.2 52.1 ( 1.5 0.07 ( 0.02 21.1 ( 12.8 1.29 45
3.4 ( 0.5 35.5 ( 3.1 0.47a 14.4 ( 10.2 1.76 77
2.5 ( 0.6 29.3 ( 3.6 0.03 ( 0.01 25.2a 4.28 641
1.9 ( 0.3 21.2 ( 3.4 0.17a 36.1a 3.09 287
Very large confidence interval.
a
c
b
d
Figure 4. Concentration profiles of hydrogen and the aromatic compound in the catalyst particles with (a) high and (b) low p-xylene concentration and with (c) high and (d) low mesitylene concentration.
The minimization was carried out by using an adaptive Levenberg-Marquardt routine, NL2SOL (Dennis et al., 1981). In the very beginning of the experiments the reactor temperature had not achieved its desired value; therefore, the first two analyses were neglected in the parameter estimation. The third analysis was used as the initial state for the reactor simulation, and only the remaining observations were included in the RSS. The mass fractions in the liquid phase were calculated from the liquid mole fractions as
wi ) xiMi/
∑i xiMi
(46)
where Mi is the molar mass of the compound i. To suppress the correlation between the parameters, the
rate equations of model I were written in the form
[
R ) Rref exp -
(
Ea 1 1 R T Tref
[
)]
×
]
cAcH(3KAcA,ref + xKHcH,ref + 1)3
cA,refcH,ref(3KAcA + xKHcH + 1)3
(47)
where the reference concentrations were cA,ref ) 3000 mol/m3 and cH,ref ) 150 mol/m3. The estimated parameters of models I (eq 21) and II (eq 29) are shown in Tables 3 and 4, respectively. Both kinetic models describe the data reasonably well (Figures 2 and 3). The beginnings of the concentration curves (Figures 2 and 3) are almost straight lines, suggesting that the reaction is nearly zero order with respect to the aromatic
4432 Ind. Eng. Chem. Res., Vol. 35, No. 12, 1996
concentration, at least at higher aromatic concentrations. Model I (eq 21) cannot, however, predict the zero order behavior for a wide concentration range, which can be seen from slightly S-shaped curves in Figure 2. The adsorption coefficient of hydrogen (KH in model I) was inadequately identified, and very different values were obtained for different aromatic compounds (Table 3), whereas approximately equal values were expected. Model II predicts the observed zero order reaction very well. The model simulations during the parameter estimation showed that the heat transfer limitation in the catalyst particles was negligible. The temperature in the center of the particles was at maximum 0.4 °C higher than the temperature of the liquid bulk. In contrast to that, the effect of the mass transfer was important. At higher aromatic concentrations (>5 wt %), the reaction rate was influenced by the hydrogen mass transfer (Figure 4a,c). After the aromatic concentration decreased significantly, the mass transfer of the aromatic compound began to affect the reaction rate (Figure 4b,d). The hydrogen mass transfer limitation is stronger in the hydrogenation of p-xylene (Figure 4a) than in the hydrogenation of mesitylene (Figure 4c). This is a consequence of the higher reaction rate of p-xylene since the diffusivity of hydrogen in p-xylene is higher than its diffusivity in mesitylene. Conclusions Hydrogenation experiments with four disubstituted benzenes and one trisubstituted benzene were carried out at two temperatures and two pressures. The reaction rate of the trisubstituted benzene (mesitylene) was observed to be lower than the reaction rate of the disubstituted benzenes (xylenes) with similar substituents (Table 2). The relative positions of the substituents appeared to affect significantly the reaction rate. The para-position was observed to be the most reactive and the ortho-position least reactive. The experiments were carried out under conditions where the internal diffusion in the catalyst particles was not negligible. Therefore, the heterogeneous reactor model was used in the estimation of the kinetic parameters. Simulations with the reactor model showed that the heat transfer resistance inside the particles was negligible, whereas the mass transfer resistances of hydrogen and the aromatic compound were important. Two mechanistic models were fitted to the experimental data. At higher aromatic concentrations, the reaction was observed to be nearly zero order with respect to the aromatic concentration (Figure 3). Model I proposed by Smeds et al. (1995) includes the competitive adsorption of the aromatic compound and hydrogen and thus cannot describe the zero order behavior for a wide concentration range (Figure 2). Model II (Temkin et al., 1989) instead predicts the zero order reaction well. Acknowledgment Financial support from Neste Oy Foundation is gratefully acknowledged. Notation a ) gas-liquid mass transfer area/reactor volume, m-1 aS ) liquid-solid mass transfer area/reactor volume, m-1 ci ) concentration of component i, mol m-3 cGi ) concentration of component i in the gas phase, mol m-3
cLi ) concentration of component i in the liquid phase, mol m-3 cp ) heat capacity of catalyst particles, kJ kg-1 K-1 Di ) diffusion coefficient of component i, m2 s-1 Di,eff ) effective diffusion coefficient of component i, m2 s-1 Ea ) activation energy, kJ mol-1 Fi ) feed flow of component i to the reactor, mol s-1 ∆HR ) reaction enthalpy, kJ mol-1 ki ) reaction rate constant, mol s-1 kg-1 kGi ) gas side mass transfer coefficient of component i, m s-1 kLi ) liquid side mass transfer coefficient of component i, m s-1 Ki ) vaporization equilibrium coefficient of component i K ) adsorption coefficient, m3 mol-1 Ki′ ) modified vaporization equilibrium coefficient of component i Kp ) constant of the P-controller, mol s-1 m-3 mcat ) catalyst mass, kg Mi ) molar mass of component i, g mol-1 nGi ) amount of component i in the gas phase, mol nLi ) amount of component i in the liquid phase, mol NGLi ) gas-liquid mass transfer flux of component i, mol m-2 s-1 NLSi ) liquid-solid mass transfer flux of component i, mol m-2 s-1 ri ) reaction rate, mol s-1 kg-1 R ) gas constant, 8.314 41 J mol-1 K-1 Rp ) catalyst particle radius, m RRMS ) residual root mean square RSS ) residual sum of squares t ) time, s T ) temperature, K VG ) gas phase volume, m3 VL ) liquid phase volume, m3 VR ) reactor volume, m3 wi ) weight fraction of component i x ) dimensionless radial coordinate in the catalyst particles xi ) mole fraction of component i Greek Letters R ) parameter in eqs 13 and 14 β ) parameter in eqs 13 and 14 γ ) number of active sites used for hydrogen molecule adsorption ) porosity of the catalyst particles λ ) thermal conductivity of catalyst particles, W m-1 K-1 F ) density of catalyst particles, kg m-3 τ ) tortuosity factor of the catalyst particles
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Received for review October 17, 1995 Revised manuscript received May 3, 1996 Accepted August 2, 1996X IE950636C
X Abstract published in Advance ACS Abstracts, November 15, 1996.
