Kinetic Modeling for Assessing the Product Distribution in Toluene

Ind. Eng. Chem. Res. 2008, 47, 1043-1050

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Kinetic Modeling for Assessing the Product Distribution in Toluene Hydrocracking on a Pt/HZSM-5 Catalyst Pedro Castan˜ o,*,† Jose´ Marıá Arandes,† Ba´ rbara Pawelec,‡ Martıń Olazar,† and Javier Bilbao† Department of Chemical Engineering, UniVersity of the Basque Country, 644-48080 Bilbao, Spain, and Institute of Catalysis and Petrochemistry, CSIC, C/Marie Curie, 2, Cantoblanco, 28049 Madrid, Spain

A kinetic model has been proposed to quantify the distribution of product lumps in the hydrocracking of toluene on a bifunctional catalyst prepared by physically mixing Pt/γ-Al2O3 with high-acidity HZSM-5 zeolite (Si/Al ) 15). The current model is based on one that was previously proposed for the hydrogenolytic cracking of methylcyclohexane, given that this is the limiting step because of the rapid hydrogenation of toluene. The experimental results used for calculating the kinetic parameters were obtained in an isothermal fixed-bed reactor under a wide range of operating conditions (250-450 °C; WHSV ) 2.85-100 h-1, τ ) 1.0-3.5 gzeolite h gtoluene-1); pressure ) 20-60 bar; molH2/moltoluene ) 19-59; conversion ) 0-100% (integral reactor). The model faithfully predicts the effects of operating conditions on the product distribution and is suitable for use in the optimization of C2+ n-alkane production. 1. Introduction Restrictions on the content of aromatics, sulfur, and nitrogen in fuels; increasing valorization of low-quality crude; and intensification in the valorization of secondary-interest refinery streams are vectors that have contributed to accelerate the implementation of hydrotreatment units in refineries, despite their high capital costs.1,2 The great versatility of hydrotreatment units makes them suitable for (i) conditioning heavy feeds prior to their valorization in FCC (fluid catalytic cracking) units, given that they improve the yield and composition of medium distillates in detriment to dry gases, cycle oil, and coke;2 (ii) adaptation (refining) of gasoline and diesel streams of FCC units, in order for their composition to fulfill market requirements;3,4 and (iii) valorization of feeds that are refractory to catalytic cracking.5-7 It is noteworthy that many units that were originally designed for mild hydrotreatment (hydrodesulfuration-hydrogenation, without hydrogenolytic cracking of C-C bonds) have undergone revamping and become hydrodesulfuration-hydrocracking units (with hydrogenolytic cracking).8 Likewise, catalysts and process technology have undergone important innovations in order to valorize feeds that are increasingly refractory and obtain fuels of higher quality.9-11 Weitkamp et al.12 justified the strategic and economic interest of the hydrotreatment of pyrolysis gasoline (PyGas) in the production of C2+ n-alkanes, which are a suitable feed for steamcrackers (whose standard feed is naphtha).13-15 PyGas is a byproduct of olefin production whose high aromatic content has meant restricted incorporation into the gasoline stream in the refinery. Ringelhan et al.16 proposed the valorization of PyGas (and, in general, of light aromatics, C9-) by the process called ARINO (aromatic ring-opening) developed by a consortium of Linde, VEBA Oil, and Su¨d-Chemie. The process is carried out in two steps: (i) hydrogenation of the aromatic ring with a catalyst of * To whom correspondence should be addressed. Currently working as a Postdoctoral Researcher in the Department of Catalysis Engineering, Delft University of Technology. E-mail: [email protected]. Tel.: +34 94 6012511. Fax: +34 94 6013500. † University of the Basque Country. ‡ Institute of Catalysis and Petrochemistry.

supported noble metal and (ii) ring-opening of the generated naphthene on an acid zeolite.13,17,18 The main advantage of the one-step process (hydrocracking) is that hydrogenation equilibrium is completely displaced because the naphthenes formed are converted by hydrogenolytic cracking.19 Furthermore, energy released in hydrogenation (exothermal) compensates cracking requirements, which is very important for isothermicity in large-scale reactors.20 Three types of reactions take place in the hydrocracking of aromatics: (i) hydrogenation-dehydrogenation, (ii) isomerization, and (iii) cracking. The latter can vary in nature: primary (ring-opening), secondary, and successive.21,22 At low temperatures, the predominant reactions are hydrogenation and skeletal isomerization, because of their lower activation energies, and as the temperature is increased, the ring-opening rate increases. The reactions that take place on the metallic function of the catalyst (of type i) are faster than those that take place on the acid function (types ii and iii), and the temperatures required for naphthenic ring scission are relatively high, above 250 °C. The latter is attributed to the fact that the intermediate carbocations that take part in the cracking have an unfavorable orbital orientation compared to the corresponding paraffins or olefins.23 In the literature, studies have been reported both on the different relevant aspects concerning the role of the active sites of the metallic and acid functions in the reaction mechanisms and on the conversion and selectivity of model aromatics transformation.21,22,24-26 Other studies have also been carried out on the effects of process conditions (temperature, space time, and hydrogen partial pressure) on the selectivity and deactivation of the catalyst by coke.27 Nevertheless, kinetic models for light aromatic hydrocracking that are useful for reactor design have not been proposed. This article focuses on this objective by proposing a kinetic model for the hydrocracking of toluene (as a model compound). The kinetic model proposed is based on the kinetic scheme for methylcyclohexane (MCH) cracking on HZSM-5, which is the determining step in hydrocracking on a bifunctional catalyst and was studied in a previous work.28 This kinetic scheme considers the effects of hydrogen and the metallic function, which is the main difference from the model proposed by Cerqueira et al.29 and subsequently tested by Caeiro et al.30 for the catalytic cracking of methylcyclohexane on an acid catalyst without hydrogen in the reaction medium.

10.1021/ie071154r CCC: $40.75 © 2008 American Chemical Society Published on Web 01/25/2008

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Ind. Eng. Chem. Res., Vol. 47, No. 4, 2008

2. Experimental Section 2.1. Catalyst. A bifunctional catalyst, Pt/z15, was obtained by physically mixing a platinum-supported catalyst (Pt/γ-Al2O3, Johnson Matthey), Pt/A, with an acid catalyst of HZSM-5 zeolite, in suitable quantities to achieve 0.5 wt % Pt/HZSM-5. Before being mixed with the acid function, the metal function was ground and sieved to obtain particle sizes in the range 0.05 < dp < 0.10 mm. The acid catalyst was prepared from an NH4-ZSM-5 zeolite supplied by Zeolyst International, with a ratio of Si/Al ) 15 (SiO2/Al2O3) 30, Na2O ) 0.05 wt %). The zeolite (25 wt %) was agglomerated with (i) bentonite (30 wt %) (montmorillonitetype clay without catalytic activity) as a binder, which confers mechanical resistance to abrasion, attrition, and erosion on the catalyst, in order to support high pressures, and (ii) an inert charge of R-Al2O3 (Merck) (45 wt %) to increase the stability of the crystalline structure and the thermal conductivity. The agglomeration provides the catalyst with the intraparticle mesopores of bentonite and alumina and also with the macropores or cages between the microparticles of individual components. The final catalyst particles were obtained by wet extrusion, drying (24 h at 120 °C), sieving in the 0.15-0.30 mm range, and final calcination (3 h at 550 °C) to achieve suitable control of dehydroxylation by forming the acid sites (Bro¨nsted and Lewis) responsible for cracking. The choice of a low Si/Al ratio in the zeolite leads to a catalyst with high acidity, as is required in severe ring-opening reactions.17,18 The metallic content of Pt/A sample was determined by means of inductively coupled plasma (ICP) spectroscopy (Perkin-Elmer Optima 3300DV). The physical properties of the catalyst functions were determined by N2 adsorption-desorption isotherms at -196 °C in a Micromeritics ASAP 2010 instrument, and the metallic dispersion was analyzed by H2 adsorption isotherm at 35 °C in a Micromeritics ASAP 2010C apparatus. The acid properties were measured by NH3 TPD in a Balzers Quastard 422 mass spectrometer on-line with a Setaram TGDSC 111 instrument at 150 °C; results obtained by this technique were consistent with those derived from thermogravimetric analysis.31-34 The concentrations of Bro¨nsted and Lewis acid sites were assessed by FTIR spectrophotometry (Nicolet 740 SX) from the areas of the peaks at 1545 and 1453 cm-1, respectively, of adsorbed pyridine at 150 °C, using the extinction coefficients given by Emeis.35 The properties of the metallic function are as follows: Pt content, 0.51 ( 0.02 wt %; BET surface area, 118 m2 g-1; pore volume, 0.26 cm3 g-1; metallic dispersion, 81%; total acidity, 160 µmolNH3 g-1. The high metallic dispersion and total acidity are noteworthy, and they will contribute to the cracking reaction. The properties of the acidic function are as follows: BET surface area, 220 m2 g-1; micropore area, 91.9 m2 g-1; micropore volume, 0.040 cm3 g-1; total acidity, 175 µmolNH3 g-1; acid strength, 152 J (mmolNH3)-1; Bro¨nsted/Lewis acid site ratio, 2.62. More detailed characterization of this catalyst can be found elsewhere.36,37 2.2. Equipment and Conditions for Reaction and Product Analysis. The reaction system has been described in detail elsewhere28 and consisted of a downward-flow fixed bed of 0.8cm internal diameter and 30.3 cm height that can operate up to 100 bar and 500 °C. Temperature was measured by a K-type thermocouple connected to a PID controller that regulated a series of resistances located around the reactor. The products leaving the reactor passed through a filter to eliminate traces of catalyst. Subsequently, the products passed through a six-port Valco valve, whereby they entered a loop and were carried by

Figure 1. Location of catalyst functions in the bed: (A) in series (hydrogenation-cracking) and (B) in the same bed (hydrocracking).

a 20 cm3 min-1 N2 stream toward the gas chromatograph. The equipment used was a Hewlett-Packard 5890 Series II instrument provided with a flame ionization detector (FID) and two columns. The first column was a semicapillary HP-1, filled with methylsilicon, of 5 m length and 0.53 mm internal diameter. The function of this precolumn was to trap the products to be analyzed. The second column was a capillary Tracer TRB-1, also filled with methylsilicon, of 60 m length, 0.20 mm internal diameter, and 0.5 µm film thickness. Given that the chromatogram obtained with an FID detector involves only the hydrocarbon stream, the conversion at the outlet was calculated by a mass balance of carbon atoms at the reactor inlet and outlet streams. Identification of reaction products was performed with a mass spectrometer (Hewlett-Packard 5989B) on-line with a gas chromatograph (Hewlett-Packard 5890 Series II Plus), equipped with a column similar to that used in product analysis. The runs were carried out under the following conditions: temperature ) 250-450 °C, WHSV ) 2.85-100 h-1 (τ ) 1.03.5 gzeolite h gtoluene-1), pressure ) 20-60 bar, molH2/moltoluene ) 19-59, conversion ) 0-100% (integral reactor). The total number of runs was 244 under 122 different conditions. All runs were repeated for statistical purposes. 3. Results 3.1. Product Distribution. Hydrocracking in a single reaction step (scheme B in Figure 1) gives way to a product distribution different from that obtained by using metallic and acid functions in series (scheme A in Figure 1). Plots a and b in Figure 2 show the product distributions at different temperatures corresponding to hydrogenation followed by cracking (two steps) and toluene hydrocracking. Conversion was determined as the fraction of toluene converted. The selectivity of each product i (Si) was calculated as the ratio between its yield (Yi) and the toluene conversion (X)

Si )

Yi X

(1)

The reaction products were grouped into the lumps methane, C2+ n-alkanes, isoalkanes, cycloalkanes, and aromatics. This grouping of components into lumps is a common strategy in the kinetic modeling of the cracking process when the aim of kinetic modeling is to use it in reactor design.38 A comparison of plots a and b in Figure 2 (corresponding to the two arrangements in Figure 1) shows that the product distributions are very similar in the 250-300 °C range. As can be observed in Figure 2a, the conversion of toluene decreases as temperature is increased above 300 °C, which is due to

Ind. Eng. Chem. Res., Vol. 47, No. 4, 2008 1045

Figure 2. Effect of temperature on conversion and selectivity: (a) scheme A (two steps in series) in Figure 2, (b) scheme B (hydrocracking) in Figure 2. Conditions: 60 bar, PH2 ) 59 bar, and WHSV ) 4 h-1.

thermodynamic limitations for hydrogenation.39 Toluene is only partially saturated on the metallic phase (because of the thermodynamic limitation), and the unconverted fraction cannot be cracked by the subsequent acid function. Consequently, the selectivity to aromatics is significant and increases exponentially with temperature. Furthermore, it cannot be ignored that part of the methylcyclohexane formed in the first bed might undergo dehydrogenation in the second bed (of zeolite) to give toluene. Because of thermodynamic restrictions, the conversion of toluene is expected to be low above 400 °C and almost null at 450 °C. Nevertheless, toluene undergoes disproportionation above 400 °C on HZSM-5 to give benzene and xylenes,40-43 which explains the high selectivity to aromatics, close to 70 wt % in Figure 2a at 450 °C. As the temperature is increased, Figure 2b (hydrocracking) shows the following evolution in the results: In the 300400 °C range, endocyclic scission of C-C bonds is favored, with formation of C2+ n-alkanes and isoalkanes, whose maximum yield is at 350 °C. Above 350 °C, the presence of cycloalkanes in the outlet stream is almost null. A C2+ n-alkane yield of 75 wt % is attained at 400 °C, and this yield is 80 wt % at 450 °C. Nevertheless, the selectivity to methane is significant above 350 °C and increases exponentially as temperature is increased, approaching 10 wt % at 450 °C. Consequently, to avoid an excessive yield of this compound, which is inert in the steamcracking unit, a reaction temperature lower than 425 °C is advisable. As an example of hydrocracking results, Figure 3 shows those for the evolution of conversion and yields with space time, obtained at 20 bar and 400 °C. Toluene hydrogenation is very rapid in the initial section of the catalytic bed, and at temperatures above 400 °C (as is the case in Figure 3), part of the toluene is disproportioned to give benzene and xylenes.44 Once cycloalkanes (cyclohexane, methylcyclohexane, and dimethylcyclohexane) are formed by hydrogenation, they undergo qualitatively the same reactions that occur in the hydrogenolytic cracking of methylcyclohexane.28 Under the conditions in Figure 3, hydrogenation and ring-opening (primary hydrogenolytic cracking) are complete for a space time of 0.1 h (no aromatics and cycloalkanes in the product stream), and so, secondary and tertiary hydrogenolytic cracking, which are responsible for the formation of linear paraffins, take place. In view of the results for different temperatures, the aromatic disappearance rate decreases as temperature is increased. This result is explained by phenomena that affect the transformation into aromatics: (1) the thermodynamic limitation of hydrogenation, which is mitigated by hydrogenolytic cracking of meth-

Figure 3. Evolution of conversion and yields with space time. Conditions: Ptoluene ) 1 bar, PH2 ) 19 bar, T ) 400 °C.

ylcyclohexane, and (2) the disproportionation of toluene to benzene and xylenes, which is favored by temperature and is responsible for the aromatics consisting of benzene, toluene, and xylenes (BTX). These phenomena are difficult to consider quantitatively in the equations of the kinetic model. Thus, the hydrogenation of the entire BTX fraction must be established, although the components of this fraction have substantially different hydrogenation rates and equilibrium constants.45,46 Furthermore, because of an equilibrium shift, the proposal of a simple kinetic model requires hydrogenation to be considered as an irreversible reaction. Figure 4 shows the evolution with space time of C2+ n-alkanes (Figure 4a) and isoalkanes (Figure 4b), at 400 °C and under different hydrogen partial pressures. The results show that, as the hydrogen partial pressure is increased, the yield of isoalkanes increases (Figure 4b) and the yield of C2+ n-alkanes decreases (Figure 4a). This result confirms the high isomerization capacity of Pt/HZSM-5 bifunctional catalyst compared to the monofunctional catalyst based on HZSM-5.47,48 Consequently, an increase in the concentration of hydrogen favors isomerization over the hydrogenolytic cracking of methylcyclohexane. The implication of this effect in kinetic modeling is that the reaction order with respect to hydrogen (R), which is 2 for the hydrogenolytic cracking of methylcyclohexane,28 must be recalculated.

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Figure 4. Effect of hydrogen partial pressure on the evolution of the yields of (a) C2+ n-alkanes and (b) isoalkanes with space time. Conditions: Ptoluene ) 1 bar, T ) 400 °C.

Figure 5. Effect of hydrogen partial pressure on the evolution of the CMR index (eq 2) with space time. Conditions: Ptoluene ) 1 bar, T ) 350 °C.

Despite the above considerations, the higher initial rate for the formation of C2+ n-alkanes (initial slope in Figure 4a) for low conversions (quasidifferential reactor) was obtained under the maximum hydrogen partial pressure. In this range, the competition between hydrogenolytic cracking and isomerization is less significant. Figure 5 shows the effect of space time and hydrogen partial pressure on the CMR (cracking mechanism ratio) index.49 The results correspond to 350 °C. The CMR index quantifies the importance of monomolecular cracking in catalytic cracking.50,51 It is calculated as the quotient between the sum of the methane and ethane yields and the yield of isobutane

CMR )

YC1 + YC2 YiC4

(2)

The results in Figure 5 show that, for low values of space time, the value of CMR initially decreases as space time is increased and subsequently increases linearly in the range where hydrogenolytic cracking (τ > 2.5 × 10-2 h) prevails. This minimum value corresponds to the maximum isoalkane (par-

ticularly isobutene) yield, as shown in Figure 4b. At low space time values, there is an incipient dealkylation (increase in methane yield), and after subsequent predominance of isomerization, hydrogenolytic cracking is enhanced. In the latter range, it is clear that an increase in hydrogen partial pressure favors the hydrogenation of carbenium ions (higher value of CMR index as seen in Figure 5). As dealkylation and protolytic cracking are unfavorable compared to the hydrogenolytic cracking and isomerization, the methane formation rate decreases significantly compared to the monofunctional catalyst,28 given that it is a consequence of these two mechanisms. This result constitutes an important advantage of the integrated process of hydrocracking over the two-step process in series. 3.2. Kinetic Model. Based on the knowledge of the effects of operating variables and corresponding determining factors, the kinetic scheme shown in Figure 6 is proposed. In this scheme, the reactions considered in the cracking of methylcyclohexane take place28 after a fast primary hydrogenation. These steps have already been considered by Cerqueira et al.29 and lately verfied by Caiero et al.30 in the catalytic cracking of methylcyclohexane without hydrogen in the reaction medium. In the kinetic modeling, the mass fraction of each lump (xi) is taken as a fitting variable. These fractions are based on the entire stream, including hydrogen. The hydrocarbon mass fractions were calculated from the GC analysis, whereas the hydrogen mass fraction was calculated from the global mass balance, according to the initial partial pressure of this compound. The differential equations that describe the formation step of each lump in the kinetic scheme shown in Figure 6 are as follows

dxA ) -k1xAx Hβ 2 dτ dxC ) k1xAx Hβ 2 - k2xCxHR 2 - k3xCxHR 2 - k6xCxHR 2 dτ dxI ) k2xCxHR 2 - k4xIxHR 2 - k5xIxHR 2 dτ dxP ) k3xCxHR 2 + k4xIxHR 2 dτ dxM ) k5xIxHR 2 + k6xCxHR 2 dτ

(3) (4) (5) (6) (7)


Figure 6. Kinetic scheme proposed for toluene hydrocracking on Pt/HZSM-5. Table 1. Kinetic Parameters for Toluene Hydrocracking Model (r ) 1) reaction

ki390°C

units

ki0

k1 k2 k3 k4 k5 k6

988 ( 55 41.7 ( 1.7 66.1 ( 2.3 2.16 ( 0.41 0.43 ( 0.25 0.55 ( 0.06

gout/(gzeolite h) (gout)2/(gH2 gzeolite h) (gout)2/(gH2 gzeolite h) (gout)2/(gH2 gzeolite h) (gout)2/(gH2 gzeolite h) (gout)2/(gH2 gzeolite h)

Ei (kJ/mol)

(4.88 ( 0.87) × 10-6 -105 ( 5 (7.78 ( 0.32) × 109 104 ( 2 (7.02 ( 0.24) × 1011 127 ( 2 (6.67 ( 0.87) × 108 107 ( 8 (5.7 ( 0.2) × 1012 166 ( 12 (2.75 ( 0.45) × 1010 135 ( 21

where the subscripts correspond to the following component lumps: C, cycloalkanes; I, isoalkanes; P, paraffins (C2+ nalkanes); M, methane; A, aromatics; and H2, hydrogen. The effect of the hydrogen concentration was considered by means of reaction orders (R and β). In the mass balance, the mass flow rate of hydrocarbons at the outlet (Qout hc ) was ) because of the assumed to be higher than that at the inlet (Qin hc incorporation of H2 into the stream. Accordingly, an experimental value was used for the parameter η

η)

Qout hc Qin hc

(8)

To close the mass balance, each kinetic constant calculated from eqs 3-7 should be divided by the parameter η. The values of the kinetic constants at the reference temperature (390 °C), the activation energies, and the reaction order R were calculated by fitting the experimental concentration data (xi) to the values of the concentrations obtained by numerical integration of eqs 3-7. The ordinary differential equations were solved using the ode113, ode15s, and ode45 codes of Matlab R14, and the values of the constants were calculated by minimizing the errors by the lsqcurvefit function, which uses the large-scale algorithm, subspace trust region method based on the interior-reflective Newton method, to solve nonlinear curve-fitting problems in the least-squares sense Nexp

SSR )

(F(k,xi) - yi)2 ∑ i)1

(9)

In preliminary calculations, the parameters R and β were fitting variables; then based on the trend of the results, values of R ) 1 and β = 1 were taken, which simplify the model and provide a physical meaning to it. Table 1 reports the values calculated for the kinetic parameters (rate constant at 390 °C, pre-exponential factor, and activation energy) of each reaction step in the kinetic scheme (Figure 6). The sum of square errors for the regression is 0.288, and the variance is 3.85 × 10-4. The negative value of the apparent activation energy of the hydrogenation reaction (k1) reported in Table 1 can be attributed to an increase in the thermodynamic limitations when the temperature is raised, which leads to the following experimental observation: Increasing the temperature caused a decrease of the hydrogenation rate (the slope of the Arrhenius plot is positive). On the other hand, it should be noted that a study

Figure 7. Partitioning diagram for the kinetic model.

was also carried out on a kinetic model that considers primary hydrogenation reversibility, which was quantified by means of an apparent equilibrium constant (there is no equilibrium constant with a real physical meaning because of disproportionation). The fitting of this model is not significantly better than that of the proposed model, although this new kinetic scheme is especially interesting under conditions in which the cycloalkane concentration is high, i.e., when the acid function is not very active, which does not occur in this study. It is particularly interesting to compare the values of the kinetic constants given in Table 1 with those found for the hydrogenation of toluene36 and the ring-opening of methylcyclohexane,28 i.e., to compare the two-step process with the single-step process for the valorization of aromatics. For toluene saturation, the hydrogenation rate increases in the integrated process because of the conversion of the products (cycloalkanes), but a direct comparison of the kinetic values is not possible because of the side reactions occurring on the bifunctional catalyst (disproportionation and cracking) and the consideration of irreversible hydrogenation during hydrocracking. For a comparison of the values in Table 1 with those for the ringopening of methylcyclohexane, one needs to refer the constants to the relative mass of zeolite in the catalyst (as in Table 1). Then, upon comparison of the hydrocracking and ring-opening, one can conclude that the former has (i) a higher formation of isoalkanes (k2), which can be formed from endocyclic scission or isomerization of n-alkanes; (ii) higher rates of isoalkane conversion to form n-alkanes (k4) and methane (k5); and (iii) a lower capacity for dealkylation of cycloalkanes (k6). Moreover, (iv) the values of R decrease from 2 to 1. All of these results can be explained by the enhanced capacity for hydrogen spillover on the bifunctional catalyst during hydrocracking, which increases the rates of isomerization and hydrogenolytic cracking. Figure 7 shows the parity diagram of the model indicating, a satisfactory fitting of the concentrations calculated using the kinetic model to the experimental data for the whole range of conditions studied. As an example of a product distribution

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simulation allows for predictions of the composition at the reactor outlet for a wide range of experimental conditions similar to those studied here. Furthermore, the effects of operating variables and the final product distribution are satisfactorily described by the model. 3. Conclusions

Figure 8. Comparison of composition values (mass fraction in the hydrocarbon stream) given by the model (lines) with experimental results (points). Conditions: T ) 400 °C, Ptoluene ) 1 bar, and PH2 ) 59 bar.

A simple kinetic model has been proposed for toluene hydrocracking that satisfactorily predicts the influence of the main process variables for a wide range of values when it is used for reactor simulation. Among the aspects that the model is able to predict, the following are notable: (i) The methane concentration is significant only for high values of space time and high temperatures. (ii) C2+ n-alkanes are formed very quickly, especially at temperatures higher than 350 °C. Above this temperature, the increase in yield is slight but more evident than in the hydrogenolytic cracking of methylcyclohexane, which is the second step in the ARINO process. (iii) The yield of isoalkanes is substantially higher than in the second step of the ARINO process. The maximum is 40 wt % compared to 28 wt % in the latter process. Furthermore, in all cases, there is a peak in the evolution of the isoalkane yield with space time and temperature. (iv) There is a rapid and irreversible formation of cycloalkanes, with a maximum (90 wt %) at low temperatures and space times. It must be noted that the kinetic model was obtained using data corresponding to a very wide range of operating conditions. Consequently, it satisfactorily takes into account the features inherent to hydrogenolytic cracking, which occurs without catalyst deactivation by coke. Likewise, the model shows that the aromatic hydrogenation-dealkylation and paraffin-cracking reactions are irrelevant. The kinetic model proposed will be a basis for the future kinetic modeling of the hydrocracking of more complex feeds, such as pyrolysis gasoline, light cycle oil (LCO) in an FCC unit, and other aromatic streams of refineries. Notation

Figure 9. Effects of temperature and space time in the reactor on the compositions (mass fractions in the hydrocarbon stream, wt %) of different lumps: (a) methane, (b) C2+ n-alkanes, (c) isoalkanes, (d) cycloalkanes. Conditions: Ptoluene ) 1 bar and PH2 ) 59 bar.

fitting, Figure 8 shows a comparison of composition results calculated using the kinetic model (lines) with experimental data (points), in the 0-0.35 h range of space times at 400 °C and PH2 ) 59 bar. The kinetic model was used to calculate contour maps for the compositions of the different lumps as a function of operating conditions. As an example, temperature and space time are the variables in Figure 9. The results shown in Figure 9 (where the darker zones correspond to the higher concentrations for each lump) and those corresponding to the other operating conditions allow for optimization of the process variables to maximize the production of C2+ n-alkanes. Figure 9 shows that the optimum for a steam cracker feedstock production corresponds to a temperature between 410 and 420 °C and to a space time higher than 0.30 h. All of these results show that the

ARINO ) aromatic ring opening BTX ) benzene, toluene, and xylenes CMR ) cracking mechanism ratio, eq 2 Ei ) activation energy of step i defined in Figure 6 FCC ) fluid catalytic cracking ki ) kinetic constant of step i defined in Figure 6 MCH ) methylcyclohexane Pi ) partial pressure of lump/compound i, bar Si ) selectivity of lump/compound i, eq 1 SSR ) sums of squared residuals for regression, eq 9 out Qin hc, Qhc ) mass flow of hydrocarbons at the reactor inlet and outlet, respectively, kg s-1 X ) conversion of toluene, % xi ) mass fraction of lump/compound i, gi gtotal-1 Yi ) yield of lump/compound i, % yi ) mass fraction of lump/compound i, gi ghc-1 WHSV ) weight hourly space velocity, h-1 Subscripts hc ) hydrocarbons A, C, I, M, P ) aromatics, cycloalkanes, isoalkanes, methane, and paraffins (C2+ n-alkanes), respectively Greek Letters R )hydrogen reaction order


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ReceiVed for reView August 23, 2007 ReVised manuscript receiVed November 6, 2007 Accepted November 7, 2007 IE071154R

Kinetic Modeling for Assessing the Product Distribution in Toluene

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