Ind. Eng. Chem. Res. 1997, 36, 2101-2109
2101
Liquid-Phase Hydrogenation Kinetics of Aromatic Hydrocarbon Mixtures S. Toppinen,† T. Salmi,*,‡ T.-K. Rantakyla1 ,§ and J. Aittamaa† Engineering, Neste Oy, P.O. Box 310, FIN-06101 Porvoo, Finland, Laboratory of Industrial Chemistry, A° bo Akademi, FIN-20500 A° bo, Finland, and Department of Process Engineering, University of Oulu, FIN-90570 Oulu, Finland
The liquid-phase hydrogenation kinetics of one multiaromatic mixture and five binary aromatic mixtures of toluene, ethylbenzene, xylenes, and mesitylene were determined in a semibatch reactor operating at a pressure of 40 bar and a temperature of 125 °C. Commercial preactivated catalyst particles of nickel-alumina were used in the experiments. In mixtures, the aromatic compounds reacted in queues so that the most reactive components started to react immediately while the least reactive components did not react until the most reactive components had been hydrogenated completely. This type of reactivity decreased with the increasing number of substituents, i.e. in the order monosubstituted > disubstituted > trisubstituted. The relative positions of the substituents affected the reaction rate so that the reactivity decreased in the order ortho > para > meta. The queue effect was described with a kinetic model based on the rapid adsorption of aromatic compounds and hydrogen and sequential addition of hydrogen to the aromatic nucleus, the first hydrogen addition step being rate determining. The kinetic parameters were estimated from a heterogeneous reactor model with nonlinear regression analysis. The kinetic model was able to describe hydrogenation kinetics of the binary aromatic mixtures. Introduction Catalytic hydrogenation of aromatic molecules to their saturated homologues is of importance for the production of intermediates for the chemical industry, e.g., the production of cyclohexane, and for the removal of aromatic compounds from solvents and fuels. The hydrogenation can in principle be carried out over several catalysts such as Pt, Pd, Rh, Ru, Co, and Ni, but Ni is the dominating catalyst on the industrial scale, mainly because of its low price. The catalytic hydrogenation is carried out in the gas phase and in the liquid phase. Gas-phase hydrogenation is applied in the synthesis of cyclohexane for the production of Nylon, whereas liquid-phase hydrogenation is the dominating process in the dearomatization units in the oil refineries. The kinetics and mechanism of the gas-phase hydrogenation of aromatic compounds have been the subject of intensive research during the years, as reviewed by Lindfors et al. (1993) and Smeds et al. (1995). The liquidphase hydrogenation of binary mixtures of alkylbenzenes on Raney nickel was investigated in the pioneering paper of Wauquier and Jungers (1957). The work concerned the hydrogenation kinetics of benzene, toluene, ethylbenzene, isopropylbenzene, xylenes, and tetraline at elevated pressures. The relative reactivities of the alkylbenzenes were determined and the effect of solvents was screened. The liquid-phase hydrogenation of the benzene ring was investigated by Temkin et al. (1989), who also proposed a reaction mechanism for the hydrogenation. The liquid-phase hydrogenation kinetics of benzene and some monosubstituted (Toppinen et al., 1996a) as well as some di- and trisubstituted (Toppinen et al., 1996b) alkylbenzenes was recently studied by our group. The experiments of our group were carried out in a laboratory scale semibatch reactor, * To whom correspondence should be addressed. Tel: +3582-2654427. Fax: +358-2-2654479. E-mail:
[email protected]. † Neste Oy. ‡ A ° bo Akademi. § University of Oulu. S0888-5885(96)00263-1 CCC: $14.00
where the liquid phase was in batch and the hydrogen pressure in the gas phase was regulated. The experiments were commenced with the pure aromatic compound (e.g., benzene), and the concentration of the compound was monitored until a complete conversion was achieved. Typical kinetic curves registered for hydrogenation of single aromatic compounds are depicted in our previous reports (Toppinen et al., 1996a,b). After an initiation period, the hydrogenation rate is virtually constant, after which the rate decreases at the end of the reaction. Numerical simulations (Toppinen et al., 1996a) based on the experimental data revealed that the kinetics were influenced by the mass transfer resistance inside the catalyst particles. In the beginning of the reaction, the hydrogen mass transfer was of crucial importance, whereas the mass transfer resistance of the aromatic compound affected the rate at the end of the reaction. The data were described with a heterogeneous model, which combined the intrinsic kinetics and the mass transfer resistances. The rate equations in the model were based on the reaction mechanisms proposed previously by Smeds et al. (1995) and Temkin et al. (1989). Both mechanisms were based on the sequential addition of hydrogen to the adsorbed aromatic molecule. The recent studies of our group (Toppinen et al., 1996a,b; Rantakyla¨ et al., 1996) concerned the hydrogenation kinetics of single aromatic molecules. In the large scale dearomatization reactors, the feed consists, however, of mixtures of unsaturated compounds. Thus the interactions between the reactant molecules may influence their reactivities. Therefore, the present work is devoted to the hydrogenation kinetics of alkylbenzene mixtures. Artificial mixtures of toluene, ethylbenzene, xylenes, and mesitylene were studied in order to determine experimentally the hydrogenation kinetics and to model the kinetics of the mixtures, particularly the interaction of competitively adsorbed aromatic compounds. © 1997 American Chemical Society
2102 Ind. Eng. Chem. Res., Vol. 36, No. 6, 1997
a
Figure 1. Simultaneous hydrogenation of four aromatic compounds at 125 °C and 40 bar.
b
Experimental Section Equipment. The liquid-phase hydrogenation studies were carried out in a stirred semibatch reactor (total volume of 1 dm3), to which hydrogen was fed continuously to maintain a constant pressure during the experiment. The reactor was heated in a three-zone electric furnace, and the temperatures of the heating zones were measured and controlled by a computer. The internal temperatures were measured with two thermocouples. The trilobe catalyst extrudates (mean particle length 3 mm, radius of the lobes 0.25 mm) were placed into the reactor in a static basket, and a magnetic propeller stirrer was regulated by a computer. A detailed description of the equipment has been given in our previous articles (Toppinen et al., 1996a,b); therefore, it is not repeated here. Catalyst and Chemicals. A commercial aluminasupported nickel catalyst (Crosfield; Ni 16.6 wt %, specific surface area 108 m2/g, mean pore volume 0.37 cm3/g, bulk density 0.86 g/cm3) was used in the experiments. Hydrogen was supplied by AGA (>99.9995%). The aromatic compounds used as mixtures were methylbenzene (toluene, >99%, J. T. Baker), ethylbenzene (>98%, Fluka), 1,2-dimethylbenzene (o-xylene, >99%, Merck-Schuchardt), 1,3-dimethylbenzene (m-xylene, >99%, J. T. Baker), 1,4-dimethylbenzene (p-xylene, >99%, Merck-Schuchardt), and 1,3,5-trimethylbenzene (mesitylene, >99%, Merck-Schuchardt). Catalyst Pretreatment. The catalyst was activated in situ by hydrogen reduction (hydrogen flow 3 dm3/min) before each experiment in order to maintain a constant activity level. The catalyst was activated in ambient pressure. During the activation the catalyst was heated to 120 °C and the temperature was maintained constant for 1 h, after which the temperature was raised to 260 °C at a rate of 80 °C/h. The hydrogen flow was continued at 260 °C for 1.5 h, after which the reactor was cooled to 80 °C under a small hydrogen flow (0.2 dm3/min). The activation procedure was automatized using the recipe and scheduling programs available in the control system package; thus the catalyst was always activated under similar conditions. Hydrogenation Experiments. The mixtures of the aromatics were prepared from commercial chemicals, which were mixed up according to the desired molar ratios. The aromatic mixture (380 mL) was loaded into the reactor through a feeding tube by using the hydrostatic pressure of the liquid. On the basis of preliminary experiments, a stirrer speed of 1800 rpm was noticed
Figure 2. Measured (symbols) and estimated (curves) aromatic concentrations in toluene-ethylbenzene experiments.
to be adequate for eliminating external mass transfer resistances. The temperature and pressure values were regulated to the desired values, and after 20 min the first liquid sample was injected automatically into a gas chromatograph (Perkin-Elmer 8500) equipped with a capillary column (30 m, J&W Scientific, DB-624) and a flame-ionization detector (FID). The column temperature was 100 °C. The duration of experiments varied from 4 to 7 h depending on the aromatic mixtures. The temperature and pressure values were registered at 1 min intervals and the sampling frequencies were typically 10-30 min. All mixture experiments were carried out at a pressure of 40 bar and a temperature of 125 °C. Results and Discussion Qualitative Results. The kinetic behaviors of the aromatics were investigated with the following mixtures: toluene-ethylbenzene, toluene-m-xylene, toluene-mesitylene, o-xylene-m-xylene, and o-xylene-pxylene and finally, with a multicomponent mixture consisting of toluene, ethylbenzene, m-xylene, and mesitylene. The mixtures were chosen to illustrate the interaction of mono- and mono-, mono- and poly-, and di- and disubstituted aromatics. The results from the multicomponent hydrogenation experiment are shown in Figure 1. The kinetic curves reveal that the behavior of the mixture deviates considerably from the behavior of the corresponding compounds, when they are hydrogenated separately. The general trend in Figure 1 is that the hydrogenation rate of a less reactive compound is considerably retarded by
Ind. Eng. Chem. Res., Vol. 36, No. 6, 1997 2103
a
a
b
b
c
c
Figure 3. Measured (symbols) and estimated (curves) aromatic concentrations in toluene-m-xylene experiments.
Figure 4. Measured (symbols) and estimated (curves) aromatic concentrations in toluene-mesitylene experiments.
the presence of a more reactive compound. For example, the hydrogenation rate of the trisubstituted compound, mesitylene, is very slow in the beginning, when unreacted mono- and disubstituted aromatics (toluene, ethylbenzene, and m-xylene) are still present in the liquid phase. The hydrogenation of mesitylene is enhanced after the mono- and disubstituted compounds have been completely hydrogenated. In a similar manner, the hydrogenation of the disubstituted compound, m-xylene, is initiated after completing the hydrogenation of the monosubstituted compounds, toluene and ethylbenzene. Toluene and ethylbenzene, on the other hand, were hydrogenated simultaneously, which is expected, since these compounds have rather equal reactivities, when they are hydrogenated separately (Toppinen et al., 1996a).
The differences in the reactivities of the model molecules, when they are hydrogenated separately, predict qualitatively the order in the mixture hydrogenation; i.e. the hydrogenation activity decreases with an increased number of substituents: monosubstituted > disubstituted > trisubstituted (Toppinen et al., 1996b). The queue effect of the mixture hydrogenation (Figure 1) can, however, not be predicted from the experiments with single components. The results depicted in Figure 1 show convincingly that a less reactive aromatic compound “waits” until the more reactive compound has been hydrogenated. In order to reveal the background of this effect the experiments with binary mixtures were carried out. The results from the mixture hydrogenation of toluene-ethylbenzene, toluene-m-xylene, toluene-mesi-
2104 Ind. Eng. Chem. Res., Vol. 36, No. 6, 1997
Figure 5. Measured (symbols) and estimated (curves) aromatic concentrations in the o-xylene-m-xylene experiment.
compound is inhibited by a less substituted aromatic compound, but not by the saturated cyclic compounds. In the hydrogenation of di- and trisubstituted aromatics (the xylenes and mesitylene) cis- and transisomers of the saturated alkylcyclohexanes are formeds the stereochemical aspects are not considered in detail here, since they are thoroughly discussed in our previous paper (Toppinen et al., 1996b). Generally, we could observe that the presence of a mixture of alkylbenzenes on the catalyst surface did not affect the stereochemical product distribution; it was the same as observed in separate hydrogenation experiments of o-xylene, mxylene, p-xylene, and mesitylene. The ratio between the cis- and trans-isomer was specific for each alkylbenzene, but virtually independent of the temperature and the hydrogen pressure. The cis-to-trans-isomer ratios decreased in the order: mesitylene (3.76) > m-xylene (3.35) > o-xylene (1.19) > p-xylene (0.89). Thus the substituents in the m-position (m-xylene and mesitylene) clearly prefer the formation of the cis-isomer. In order to describe the kinetic effects quantitatively, the reaction kinetics and the dynamic behavior of the reactor are modeled, and the rate and adsorption parameters are determined from the experimental data. Modeling of the Hydrogenation Kinetics. The kinetic model is based on the following surface reaction mechanism proposed earlier by Smeds et al. (1995) for the gas-phase hydrogenation of aromatics. The mechanism consists of aromatics and hydrogen adsorption steps, consecutive hydrogen addition steps, and product desorption steps. The mechanism can be summarized as follows: Ki
Ai + * \ y z Ai* Figure 6. Measured (symbols) and estimated (curves) aromatic concentrations in the o-xylene-p-xylene experiment.
tylene, o-xylene-m-xylene, and o-xylene-p-xylene with different concentration ratios are shown in Figures 2-6. The primary observation is that the concentration ratio of the reactants is of minor importance for this kinetics. An example is given in Figure 4 for the pair toluenemesitylene. The monosubstituted compound, toluene, is always hydrogenated much more rapidly, even though it is present as a minority ingredient (30 wt %) in the mixture. The binary mixtures confirm the fundamental observation made with the multicomponent mixture: the less reactive, more substituted compound waits until the more reactive, less substituted compound is completely hydrogenated. The compounds whose reactivities are close, e.g. toluene and ethylbenzene, are, on the other hand, hydrogenated simultaneously, as shown in Figure 2. The reactivity of the less reactive compound can be still more suppressed by an excess of the more reactive compound, as can be seen from the results obtained with the mixture, whose molar ratio tolueneto-mesitylene was 3:1 (Figure 4a). The chemical reason of the interesting queue effect might be the difference in the adsorptivities of the aromatic compounds. For instance, the experiments with pure mesitylene (Toppinen et al., 1996b) showed that the reaction starts immediately. However, in the presence of toluene, mesitylene is not able to react. This indicates that mesitylene adsorption is prevented by toluene, but not by the reaction product, methylcyclohexane. Thus we remain with the important conclusion: the adsorption of a more substituted aromatic
KH
H2 + 2• \ y z 2H• ki
Ai* + 2H• 98 AH2i* + 2• rapid
AH2i* + 2H• 98 AH4i* + 2• rapid
AH4i* + 2H• 98 AH6i* + 2• rapid
AH6i* 98 AH6i + *
(I) (II) (III) (IV) (V) (VI)
where * and • denote the active sites for hydrocarbon and hydrogen adsorption, respectively. Ai is the aromatic reactant and AH6i is the saturated product. For the case of competitive adsorption, the sites * and • coincide. The true character of hydrocarbon adsorption is probably between competitive and noncompetitive adsorption, because the larger hydrocarbon molecules cannot compete away all hydrogen; at a complete coverage of the hydrocarbons, some interstitial sites likely remain accessible for hydrogen adsorption. The adsorbed intermediates (AH2i*, AH4i*, and AH6i*) are not discussed in detail here, since an investigation of their chemical nature is beyond the available experimental technique. For gas phase hydrogenation, it has, however, been proposed that the partially hydrogenated intermediates AHi* and AH4i* have an aromatic character (Smeds et al., 1995); the cyclohexadiene type intermediates are excluded because of thermodynamic obstacles. In order to obtain rate equations from the mechanism (I)-(VI) some fundamental assumptions are introduced. The adsorption steps for hydrogen and aromatics are
Ind. Eng. Chem. Res., Vol. 36, No. 6, 1997 2105
presumed to be rapid so that the quasi-equilibrium hypothesis can be applied to these steps. The first hydrogen addition step, III, is presumed to be rate determining, whereas the subsequent reaction and desorption steps are assumed to be rapid. The overall reaction is considered to be irreversible, which confirms the experimental observations: a complete conversion of the aromatic molecules was observed in all experiments (Figures 1-6). Quantitatively, the velocity of the rate determining step (III) is given by
Ri ) kiθAiθH2
(1)
The application of the quasi-equilibrium hypothesis on the adsorption steps (I-II) gives the coverages of Ai* and hydrogen:
θAi ) KicAiθv
(2)
θH ) (KHcH2)1/2θs
(3)
Reactor Model. The three-phase semibatch reactor is described with a heterogeneous model, which accounts for the mass transfer resistance at the gas-liquid interface as well as the mass and heat transfer resistances in the porous catalyst particles. The bulk phases in the reactor and the catalyst particles are principally nonisothermal. Nonisothermicity of the particles is described with an energy balance, whereas the temperature fluctuations in the bulk phase are included directly from the experimental temperatures, thus avoiding the use of the bulk-phase energy balance. The diffusional fluxes at the gas-liquid interface and inside the catalyst particle are described with the Fickian approximation. Hydrogen was continuously fed into the reactor during the experiments; thus the hydrogen feed is modeled with a proportional (P) controller. The reactor is assumed to be vigorously stirred, so that concentration and temperature gradients in the bulk phases are negligible. Based on these concepts, the mass balances of component i for the gas (G) and liquid (L) phases can be written as follows:
where θv and θs denote the vacant sites for hydrocarbon and hydrogen adsorption, respectively. For the case of noncompetitive adsorption separate site balances are written for hydrogen and hydrocarbons. The site balance for the aromatic hydrocarbons is
dnGi ) -VRNGLia + Fi dt
(10)
dnLi ) VRNGLia - VRNLSias dt
(11)
N
θA + θv ) 1 ∑ j)1
(4)
j
It should be noticed that the surface coverages of the partially hydrogenated intermediates and the product molecules are discarded in eq 4 above; the surface is assumed to be dominated by the aromatics (toluene, ethylbenzene, xylenes, etc.). This is an implication of the hypothesis of rapid and irreversible steps IV-VI. After the quasi-equilibria are inserted in the site balance, an expression for the vacant sites is obtained:
where NGLi and NLSi are the fluxes from the gas bulk to the liquid bulk and from the liquid bulk into the catalyst particles, respectively. Fi denotes the feedsthis term is of importance for hydrogen only. The fluxes NGLi are obtained from the two-film theory. For gas and liquid films in series the Fickian law gives for NGLia:
cGi - cLi K′i NGLia ) 1 1 + kGiaK′i kLia
N
KjcA + 1)-1 ∑ j)1
θv ) (
j
(5)
For hydrogen the site balance is written as
θH + θs ) 1
θs ) [(KHcH2)1/2 + 1]-1
(7)
The expressions (2) and (3) for the adsorbed reactants are now inserted in the rate equation (1), which becomes
Ri ) kiKiKHcAicH2θvθs2
kiKiKHcAicH2 N
KjcA + 1)[(KHcH )1/2 + 1]2 ∑ j)1
(
j
2
NLSiVRas )
2
FpRp
Di,eff
( ) dci dx
(13)
x)1
(
)
∂ci Di,eff ∂2ci 1 ∂ci Fp ) + + r ∂t R 2 ∂x2 x ∂x p i p p λ
∂T )
(9)
2mcat
where (dci/dx)x)1 is the concentration gradient at the outer surface of the catalyst particle. The trilobe structure of the catalyst particles was approximated with an infinite cylinder. The concentration gradients at the surface of the particle are therefore obtained from the mass and energy balances for the particles:
(8)
By inserting the expressions (5) and (7) in (8) the following rate equation is obtained:
Ri )
For the flux from the liquid bulk into the catalyst particles, the following expression is obtained from the fundamental geometric consideration:
(6)
After the quasi-equilibrium expression for hydrogen, eq 3, is inserted in eq 6, the fraction of vacant hydrogen adsorption sites is obtained:
(12)
∂t
(
∂2T
FpcpRp2 ∂x2
)
1 ∂T + x ∂x
1 +
(14)
N
∑(-∆HR )Rj
cpj)1
j
(15)
The following initial and boundary conditions are applied:
2106 Ind. Eng. Chem. Res., Vol. 36, No. 6, 1997
nGi ) n0Gi
t)0
(i)
nLi ) n0Li
t)0
(ii)
ci ) c0i
t ) 0, x ∈ [0, 1]
(iii)
T ) T0
t ) 0, x ∈ [0, 1]
(iv)
ci ) cLi
t g 0, x ) 1
(v)
T ) TL
t g 0, x ) 1
(vi)
∂ci/∂x ) 0
t g 0, x ) 0
(vii)
∂T/∂x ) 0
t g 0, x ) 0
(viii)
The ODE-solution subroutine LSODE (Hindmarsh, 1983) was used in the numerical simulations. Parameter Estimation Procedure. The objective function (Q) for the parameter estimation was defined as follows:
(16)
The boundary conditions (16 vii) and (16 viii) arise from the particle symmetry. The amounts of substance (n) and the concentrations (c) in the liquid phase are related by
nLi
ci )
FL
N
(17)
nLjMj ∑ j)1
p D τp i
(18)
where p/τp is the particle porosity-tortuosity ratio and Di is the molecular diffusion coefficient of compound i. The molecular diffusion coefficient is obtained from Wilke’s approximation (Wilke, 1950; Reid et al., 1987), where the diffusion coefficients at the infinite dilution are calculated from the Wilke-Chang equation (Wilke and Chang, 1955).
Di )
1 - xi N
∑ j)1
(19)
wAi )
xAiMAi
nLAi
(23)
∑nLi
i.e. from amounts in the liquid bulk. The estimated weight fractions were obtained from the solution of the mathematical models, i.e. eqs 10, 11, 14, and 15, and from eq 22 for the different sampling times. The objective function was minimized with the Levenberg-Marquardt method (Marquardt, 1963) using the subroutine NL2SOL (Dennis et al., 1981) coupled with the ODE-solver LSODE. Parameter Estimation Results. In the experiments with single aromatic compounds the observed reaction rates were virtually constant, i.e. the apparent order of reaction with respect to the aromatic concentration was nearly zero, down to very low aromatic concentrations. This kind of behavior can be explained by almost total coverage of the catalyst active sites by aromatic molecules. In our kinetic model the high coverage is described with large values for the adsorption coefficients of the aromatics. At higher aromatic concentrations: N
N
KjcA + 1 ≈ ∑KjcA ∑ j)1 j)1 j
cGtot cLtot
(22)
∑xjMj
where xAi is the mole fraction calculated from
j*1
The gas-liquid equilibrium ratio (K′i) in eq 12 is calculated from
(20)
The thermodynamic equilibrium constants (Ki) are estimated from the Soave-Redlich-Kwong equation of state (Graboski and Daubert, 1978). The mathematical model consists of ordinary differential equations (ODEs) (10) and (11) and parabolic partial differential equations (PDEs) (14) and (15). The PDEs were converted to ODEs by discretizing the spatial coordinate (x) with central difference formulas. Usually, a five-point central difference formula was sufficient to achieve the desired accuracy in the approximation. Thus the coupled system of PDEs and ODEs was transformed to ODEs, to an initial value problem with respect to time. The ODEs were solved with the backward difference method (Henrici, 1962; Gear, 1971) during the estimation of kinetic parameters.
(21)
where i refers to the aromatic compound and t to the sampling time, w and w ˆ denote the experimentally recorded and the estimated weight fractions, respectively. The experimental weight fractions (wAi,t) are obtained from the fundamental relationship
(xj/D0ij)
K′i ) Ki
∑i ∑t (wAi,t - wˆ Ai,t)2
xAi )
where FL is the density of the reaction mixture. The mathematical model includes several physical and transport parameters, whose numerical values are estimated from appropriate correlations. The effective diffusion coefficients are estimated from the simple equation
Di,eff )
Q)
(24)
j
and thereby the rate equation (9) can be written as
Ri ≈
(
kiKHcAicH2 N
cAi +
Kj
cA ∑ j)1K j*1
i
)
j
[(KHcH2)
(25) 1/2
2
+ 1]
Thus, the absolute values of the adsorption coefficients have no effect on the reaction rate at higher aromatic concentrations. The rate equation (25) contains only (N - 1) ratios of the adsorption coefficients. Since most of our data were from higher aromatic concentrations, all the adsorption coefficients could not be estimated reliably. On the other hand, the new form of the rate expression (eq 25) is mathematically unsatisfactory since it cannot describe the reaction rate when the aromatic concentrations become zero. Therefore, toluene was chosen as a reference compound and a large constant value (1.0 m3/mol) was given to its adsorption
Ind. Eng. Chem. Res., Vol. 36, No. 6, 1997 2107 Table 1. Initial Compositions of the Aromatic Mixtures in the Experiments
Table 4. Estimated Parameters for the Hydrogenation of the Toluene-Mesitylene Mixture
initial composition (wt %) ethylomptoluene benzene xylene xylene xylene mesitylene other figure 25.15 74.12 27.25 74.83 49.65 26.08 73.55 49.89 26.69
24.40 25.85 72.55
24.63
25.45
0.37 0.03 0.20 0.10 0.12 1.07 0.03 0.20 0.19 0.90 0.24
25.07 50.23 72.85 26.42 49.91 73.12 49.57 49.35
49.53 50.41
1 2a 2b 3a 3b 3c 4a 4b 4c 5 6
Table 2. Estimated Parameters for the Hydrogenation of the Toluene-Ethylbenzene Mixture parameter
toluene
ethylbenzene
ki, mol/(s kg) Ea, kJ/mol Ki, m3/mol KH, m3/mol RRMSd Q
0.037 ( 0.012a 11.4 ( 4.8 (1.00)b 0.593c 4.61 3167
0.027 ( 0.010 15.7 ( 5.9 1.15 ( 0.33
toluene
mesitylene
0.055 ( 0.011 13.9 ( 2.7 (1.00) 0.099 ( 0.063 2.65 1515
0.017 ( 0.003 36.7 ( 2.3 0.07 ( 0.03
Table 5. Estimated Parameters for the Hydrogenation of the Xylene Mixtures parameter
o-xylene
m-xylene
p-xylene
ki, mol/(s kg) Ea, kJ/mol Ki, m3/mol KH, m3/mol RRMS Q
0.019 ( 0.005 41.4 ( 3.6 0.70 ( 0.17 0.052 ( 0.040 4.21 3105
0.021 ( 0.006 51.2 ( 4.4 (0.31)
0.040 ( 0.012 33.1 ( 6.6 0.49 ( 0.19
Table 6. Relative Reactivities, Rate Parameters, and Adsorption Coefficients S′Ti
kT/ki
KT/Ki
this this this work W & Ja work W & J work W & J
a 95% confidence interval. b Constant value used in the estimation. c Very large confidence interval. d Residual root mean square.
Table 3. Estimated Parameters for the Hydrogenation of the Toluene-m-Xylene Mixture toluene
m-xylene
ki, mol/(s kg) Ea, kJ/mol Ki, m3/mol KH, m3/mol RRMS Q
0.062 ( 0.015 15.8 ( 3.2 (1.00) 0.067 ( 0.047 3.05 1418
0.019 ( 0.004 50.2 ( 2.9 0.31 ( 0.07
coefficient. Adsorption coefficients of the other aromatic compounds were estimated using eq 9 for the reaction rate. The temperature dependences of the adsorption coefficients were presumed to be negligible and the rate constants ki were assumed to follow the Arrhenius’ law:
)]
Ea 1 1 ki ) ki(Tref) exp R T Tref
toluene-ethylbenzene toluene-o-xylene toluene-m-xylene toluene-p-xylene toluene-mesitylene a
parameter
[ (
parameter ki, mol/(s kg) Ea, kJ/mol Ki, m3/mol KH, m3/mol RRMS Q
(26)
where a temperature of 100 °C was used as the reference temperature Tref. The parameters of the kinetic model were estimated separately for toluene-ethylbenzene (Table 2; Figure 2), toluene-m-xylene (Table 3; Figure 3), and toluenemesitylene (Table 4; Figure 4) mixtures and simultaneously for o-xylene-m-xylene and o-xylene-p-xylene mixtures (Table 5; Figures 5 and 6). In each fit one of the adsorption coefficients was fixed and the other was estimated. The estimated value of the m-xylene adsorption coefficient (Table 3) was used as a constant in the fit with the xylene data (Table 5). The initial compositions of the reaction mixtures are given in Table 1. The data from the single aromatic experiments of the involved components (Toppinen et al., 1996a,b) were included in the fits. The systematic deviations between the experimental and estimated concentrations, especially in tolueneethylbenzene and xylene-xylene data, probably result from slight random variations in the catalyst activity level.
1.2 5.1 10.5 3.5 46
1.2 3.5 5.3 4.2
1.4 3.7 3.3 1.7 3.2
1.7 4.9 3.7 2.9
0.87 1.4 3.2 2.8 14.3
0.71 0.72 1.4 1.4
Wauquier and Jungers (1957).
The estimated rate and adsorption parameters in Tables 2-5 enable the calculation of the relative reactivities of the aromatic compounds. The division of the rate expression (9) for two aromatic compounds (i, j) gives the relative reactivity. When toluene (T) is used as a reference compound, we obtain the relative reactivity
STi )
kTKTcT kiKici
(27)
Furthermore, the ratio of the rate and equilibrium constants,
S′Ti )
kTKT kiKi
(28)
as such, is a measure of the relative reactivity and can be used in comparisons with literature data. In Table 6 we present the relative reactivities as well as the ratios of the rate parameters (R′iT ) ki/kT) determined in this work of Wauquier and Jungers (1957). In spite of different catalysts (Ni/Al2O3 versus Raney Ni) the trends are remarkably similar: we obtained the relative reactivities in decreasing order: toluene > ethylbenzene g p-xylene > o-xylene > mxylene . mesitylene, whereas Wauquier and Jungers (1957) report the reactivity orders toluene > ethylbenzene > o-xylene g p-xylene > m-xylene. The difference between the reactivities of o- and p-xylene is, however, not large, as illustrated in Figure 6. The comparison of the ratios of the rate parameters reveals that these parameters decrease in the following order toluene > ethylbenzene > p-xylene > m-xylene ≈ mesitylene > o-xylene. Wauquier and Jungers (1957) obtained the ratios of the rate constants at 170 °C: toluene > ethylbenzene > p-xylene > m-xylene >
2108 Ind. Eng. Chem. Res., Vol. 36, No. 6, 1997
a
b
Figure 7. Concentration profiles of (a) hydrogen and (b) toluene and mesitylene in the catalyst particles (experiment in Figure 4a; tmcat/VL ) 5.4 min g/mL).
o-xylene (mesitylene was not studied by them). Thus the observations are in complete agreement. Furthermore, the adsorption coefficients decrease in the order: ethylbenzene > toluene > o-xylene > pxylene > m-xylene > mesitylene. Wauquier and Jungers (1957) report the following order: ethylbenzene ≈ o-xylene > toluene > p-xylene ≈ m-xylene. For pure component hydrogenation we obtained previously (Toppinen et al., 1996b) the decreasing xylene reactivities in the order para > meta > ortho. This seems, in fact, to correspond with the ratios of the rate coefficients. However, because the adsorption affinity of m-xylene is lower than that of p- and o-xylenes, it is partly competed away in the hydrogenation of binary mixtures, and its observed reactivity is lower than the reactivities of o- and p-xylenes. Mesitylene resembles structurally m-xylene, having substituents in the meta-position. Thus mesitylene has the lowest adsorption coefficient and the lowest reactivity. 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 deg 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 only by the hydrogen mass transfer but after the aromatic concentrations decreased significantly, the mass transfer of the aromatic compounds began to affect the reaction rate. Typical concentration profiles of hydrogen and aromatic compounds are shown in Figure 7. Conclusions The hydrogenation of multiaromatic mixtures showed a queue behavior (Figure 1) where only the most
reactive aromatic components started to react immediately, whereas the less reactive components “waited” until the more reactive components were hydrogenated completely. The reactivities of alkylbenzenes decreased in the order monosubstituted > disubstituted > trisubstituted, and the less reactive di- and trisubstituted compounds were those having a substituent in the metaposition. This behavior was assumed to be a result of different adsorptivities of the components, and a kinetic model based on the competitive adsorption of the aromatic compounds was derived. A 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 compounds were important. The order of reaction with respect to the aromatic concentration was near zero. Since most of the data were from relatively high aromatic concentrations, an accurate estimation of the adsorption coefficients was not possible. The ratios between the adsorption coefficients of different aromatics could, however, be estimated. The comparatively simple kinetic model described the data well (Figures 3-6). The estimation results showed a strong dependence of the adsorption coefficient on the number of substituents in the benzene ring (Tables 2-4), as could be expected on the basis of the qualitative observations. The adsorptivity decreased with the increasing number of substituents. Also, the relative positions of the substituents seemed to affect the adsorptivity so that the adsorptivity decreases in the order ortho > para > meta (Table 5). This is reasonable because the geometric shape of the ortho form of a disubstituted benzene is close to the shape of a monosubstituted benzene. 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′i ) modified vaporization equilibrium coefficient of component i 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
Ind. Eng. Chem. Res., Vol. 36, No. 6, 1997 2109 N ) number of aromatic compounds in the mixture 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 Q ) residual sum of squares ri ) reaction rate, mol s-1 kg-1 R ) gas constant, 8.314 41 J mol-1 K-1 Ri ) rate of reaction, mol s-1 kg-1 Rp ) catalyst particle radius, m RRMS ) residual root mean square t ) time, s T ) temperature, K 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 p ) porosity of the catalyst particles θi ) fractional coverage of component i λ ) thermal conductivity of catalyst particles, W m-1 K-1 Fp ) density of catalyst particles, kg m-3 τp ) tortuosity factor of the catalyst particles
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Received for review May 10, 1996 Revised manuscript received November 29, 1996 Accepted February 11, 1997X IE960263V
X Abstract published in Advance ACS Abstracts, April 1, 1997.
