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Brazilian Journal of Chemical Engineering

Print version ISSN 0104-6632On-line version ISSN 1678-4383

Braz. J. Chem. Eng. vol.33 no.4 São Paulo Oct./Dec. 2016 



R. Thakur1  * 

S. Barman1 

R. Kumar Gupta1 

1Department of Chemical Engineering, Thapar University, Patiala-147004, Punjab, India. Phone: + 91-175-2393437; Fax: + 91-175-239300. E-mail:


In the present study, transalkylation of 1,4-diispropylbenzene (DIPB) with benzene in the presence of modified beta zeolite was performed to produce cumene in a fixed bed reactor. Beta zeolite was exchanged with cerium in order to modify its catalytic activity. Activity of the modified catalyst was evaluated in the range of temperature 493K-593K, space time 4.2 kg h/kmol-9.03 kg h/k mol and benzene/1,4-DIPB molar ratio 1-15 to maximize the reactant conversion and selectivity of cumene. The activity and selectivity of the modified catalyst was found to increase with increase in cerium loading. Maximum selectivity of cumene (83.82%) was achieved at 573 K, benzene/1,4-DIPB 5:1 at one atmosphere pressure. A suitable kinetic model for this reaction was proposed from the product distribution pattern following the Langmuir-Hinshelwood approach. Applying non-linear regression, the model parameters were estimated. The activation energy for the transalkylation reaction was found to be 116.53 kJ/mol.

Keywords: Cerium; Beta zeolite; Kinetic study; Transalkylation; DIPB; Benzene


Cumene is a colorless liquid, also known as cumol or isopropyl benzene having the boiling-range motor fuel of high antiknock value. It is of industrial demand for the production of high molecular weight hydrocarbons such as cymene and polyalkylated benzene. The main end uses for cumene are for the production of phenolic resins, bisphenol A, and caprolactam. However, 5-10 wt% diisopropylbenzene (DIPB) isomers are produced as low value byproduct during the isopropylation of benzene to cumene (Leu et al., 1990; Sridevi et al., 2001; Reddy et al., 1993). The by-products, DIPB isomers, can be recycled for cumene production, making this process more economical. With the liquid catalysts, there are inherent problems of product separation, recycling and corrosiveness (Maity and Pradhan, 2006; Barman et al., 2005; Ercan et al., 1998). In that respect, zeolites can exhibit acidities close to those of traditional mineral acid solutions and hence proved to be better catalyst (Best and Wojciechowski, 1978; Slaugh, 1983; Bakas and Barger, 1989). Moreover, the number and strength of acid sites in zeolite can be changed to a great extent by exchanging its H+/Na+ ions with rare earth cations in the zeolite framework. A comparative study was carried out on transalkylation of DIPB with benzene over Y, beta and mordenite with different Si/Al molar ratios in supercritical CO2 and liquid phase (Sotelo et al., 2006). The influence of Si/Al ratio on the activity of catalyst was explained in terms of cumene selectivity and yield considering the competitive isomerization and by-product formation. The use of supercritical CO2 did not show superior catalytic transalkylation activity for the Y zeolite. In Mobil Oil Corp., USA, production of cumene was carried out by introducing the feed to a transalkylating zone over beta zeolite/alumina and then feeding to an alkylating zone where MCM-22/alumina catalyst was used (Collins et al., 1999). Transalkylation of DIPB has also been carried out over large pore zeolites, which proved to be very active catalysts (Pradhan and Rao, 1993). In another process, DIPBs were recycled for transalkylation in the reactor containing a single catalyst bed of beta catalyst. The combined alkylation and transalkylation was performed for alkyl aromatic production to evaluate the performance of different catalysts like MCM-22 and beta zeolite based on their Si/Al ratio, selectivity, and pore size for liquid phase production of cumene (Perego and Ingallina, 2004). The catalysts such as zeolite X, MCM-22, MCM-49, PSH-3, SSZ-25, zeolite Y, beta zeolite (Yeh et al., 2008; Barger et al., 1989; Huang et al., 1997) were used in transalkylation reaction. These studies show that choice of catalyst, its Si/Al ratio and the acidity of the catalyst highly affect the process.

Kinetics of transalkylation of diisopropylbenzene were studied over Ca modified YH zeolite catalyst which proved to be a good active catalyst (Grigore et al., 2001). Cumene synthesis over beta zeolite has been reported in the literature (Bellussi et al., 1995; Perego et al., 1996; Smirnov et al., 1997; Halgeri and Das, 1999). Therefore, further investigation was necessary to carry out transalkylation of DIPB with benzene over the modified beta zeolite to obtain higher cumene selectivity and reactant conversion. Replacement of sodium ions in zeolites with polyvalent cations like rare earth metals (La, Ce, etc.) has been reported to produce materials of superior catalytic activity (Venuto et al., 1966; Rabo et al., 1968; Hunter and Scherzer, 1971). However, very scarce literature is available on the use of rare earth metal modified beta zeolite for cumene synthesis. It was, therefore, thought desirable to investigate the kinetics of this commercially important reaction over zeolite H-beta modified by exchanging H+ ions with cerium ions. A further objective of this study was to develop a suitable kinetic model for the synthesis reactions.



Beta zeolite, 1.5 mm extrudates, used in the present study, was obtained from Sud chemie, Vadodra, India. Ceric ammonium nitrate (99% pure) was procured from CDH chemicals, India. Benzene and 1,4-DIPB of analytical reagent grade ((>99% pure) were obtained from Sigma Aldrich Pvt. Ltd., India. Nitrogen gas (grade -I, 99.999% pure) was obtained from Sigma gases and services (India).

Catalyst Preparation

The commercially available H-beta zeolite containing H+ ions was modified with Ce4+ ions. At first, the zeolite extrudates were calcined for 3 h at 623 K. Calcined zeolite was then refluxed with the required percentage of ceric ammonium nitrate solution at 363 K for 24 h, thereby modifying H-beta zeolite into the Ce-beta form. The catalyst particles were then filtered and washed several times with deionized water and then dried at 393 K for 14 h. Finally, they were calcined for 4 h at 723 K to remove the excess ions. The cerium-exchanged zeolite was characterized by TPD, XRD and FTIR. Beta zeolite treated with 4%, 6%, 8%, and 10% cerium ammonium nitrate solution (CeB4, CeB6, CeB8, and CeB10) was used for the present study.

Determination of Cerium in the Exchanged Catalysts

The amount of cerium ions exchanged with the H+ ions was calculated analytically (Krishnan et al., 2002). Freshly calcined cerium modified beta zeolite was taken in a flask and digested for 1 h in concentrated HCl. The digested catalyst was diluted with distilled water and filtered. The filtrate was transferred to a beaker and its volume was made up to about 250 ml by adding distilled water. 50 ml of saturated oxalic acid solution was mixed with this solution, which produced a white precipitate of cerium oxalate. The precipitate was then filtered using a Whatman no.40 ashless filter paper and washed with distilled water. The filter paper was ignited in a previously weighed silica crucible at 1173?10 K to a constant weight. On heating, cerium oxalate was converted to cerium oxide. From the weight of cerium oxide the percentage of cerium was then calculated. CeB4, CeB6, CeB8 and CeB10 were found to have been loaded with 2.87%, 4.46 wt%, 6.64 wt% and 8.34 wt% of cerium respectively.

Experimental Setup for Transalkylation Reactions

Vapor phase transalkylation reaction was carried out in a fixed-bed, continuous down-flow, stainless steel (SS 316) reactor. The reaction conditions were maintained at atmospheric pressure. A preheater was fitted with the reactor in the upstream and a condenser in the downstream. A thermowell extending from the top of the reactor to the centre of the bed was used to measure the temperature of the reactor. Typically, 0.002 kg of the catalyst supported on a wire mesh was loaded into the reactor. Before conducting the experiments, catalyst activation was done at a temperature 100 K higher than the reaction temperature (maintained according to reaction conditions), for 3 h under the atmosphere of nitrogen. A dosing pump was used to introduce the reactant feed mixture into the reactor. Nitrogen gas was flown through the reactor at the rate of 0.565 L/h to activate the catalyst before experimental runs. However during all experimental runs, the nitrogen to feed flow rate ratio was kept constant at 0.2. The reactants were vaporized in the preheater, which is maintained at a temperature 30 K lower than the reaction temperature. The vaporised reactant feed mixture passes through the catalyst bed in the reactor at proper reaction conditions. The product vapors, along with the unreacted reactants, were condensed in the condenser (277 K-279 K). The samples were collected and analyzed in a gas chromatograph (Bruker, Model: 436 GC Scion) using a fused silica capillary column having 10 m × 0.53 mm × 1.5 µm dimensions. The sample was introduced through a micro syringe into the injector port of the GC. The temperature of the injector was set at 493 K during the analyses. The column temperature was initially set at 323K, and then increased to 523 K at a rate of 10 K/min. The flow rate of carrier gas (nitrogen) was maintained at 1.5 L/h. A Flame Ionisation Detector was used at 553 K to detect the products. Peaks were identified by retention time matching with known standards. Various products like aliphatics (propene), benzene, toluene, xylene (C8), cumene, cymene (C10), isomers of DIPB were found. The selectivity of cumene, 1,3 DIPB and conversion of 1,4 DIPB were calculated as:

The mechanism of transalkylation of 1,4 DIPB with benzene is shown in Figure 1.

Figure 1 Mechanism of transalkylation of 1,4-DIPB with benzene. 


Characterization of the Modified Catalyst

The powder X-ray diffraction patterns of the zeolite samples were recorded on a Bruker Angle X-Ray Scattering Diffractometer D8, Germany with 2θ value in the range of 50-600 at a scanning speed of 20 (2θ) per minute. The diffractometer was equipped with a Ni-filtered Cu K-α radiation source. Sample preparation for the X-ray analysis involved gentle grinding of the solid into a fine powder. Approximately 0.1-0.2 g of the sample was taken into the sample holder with light compression to make it flat and tight. Ammonia Temperature Program Desorption of the parent and modified catalysts were performed in a CHEM-BET 3000 instrument (Quantachrome). Catalyst sample (0.1 g) was first degassed at 723 K for 1 h with nitrogen followed by cooling to 273 K temperature. Nitrogen-ammonia gas (1 mol%) mixture, was then passed through the sample for 1 h. The catalyst sample was heated to 372 K until the steady state was attained; thereafter the temperature of the sample was raised up to 1173 K at a heating rate of 10 K/min. The desorbed ammonia was detected by a Thermal Conductivity Detector analyzer. Studies of FTIR were done using an Agilant Cary-660 spectrophotometer in the range 400 cm-1-4000 cm-1. Samples were analysed by preparing KBr pellets. Sample was first ground thoroughly and then it was mixed with KBr salt and again ground well to make a fine powder, which was then pressed to make a pellet. It was then placed in the FTIR sample holder and 32 scans were taken for each sample.

Both modified and unmodified beta zeolites were characterized by XRD and are shown in the Fig. 2. Typical diffraction characteristics of BEA topology are observed for both the zeolites indicating the framework structure is well preserved after cerium exchange. The CeB10 diffractograms exhibited a high intensity reflection at 2θ = 7.80 and 210-220 which is a typical characteristic peak of beta zeolite (Zhang et al., 2014). This phenomenon indicates that the crystal structure of beta zeolite was not changed by metal exchange. The crystallinity of the modified sample was measured from the intensity of the peaks. At low loading of metal ions into the zeolite framework, no diffraction lines corresponding to metal species were observed as at low loading, the cerium ions replace the H+ ions (Tang et al., 2014) and are highly dispersed in the zeolite framework (Garcia et al., 2011). However, on increasing the metal loading, additional peaks of metal oxides present on the catalyst surface appear in the XRD pattern (Siregar and Amin, 2006, Amin and Anggoro, 2003; Garcia et al., 2011). At high loading of cerium (10 wt%), additional peaks related to cerium oxide (2θ = 13.560 and 46.50) were found in zeolite XRD pattern (Thakur et al., 2015). Similar observation was reported for high loading of cerium (16 wt%) into the beta zeolite framework (Wang and Zou, 2003).

Figure 2 XRD of beta and cerium modified beta zeolite. 

The temperature programmed desorption profile of ammonia was studied with H-beta and CeB10 beta zeolites and is shown in Fig. 3. The profiles show that the catalysts contain mainly two types of acid sites of different strengths. The low temperature peak is due to ammonia desorption from the weak acid sites, whereas, the peak at high temperature is due to the desorption of strongly adsorbed ammonia from strong acid sites. TPD profiles show that, in the case of H-beta zeolite, only weak acid sites are present. In the case of Ce-beta zeolite there is an additional peak in the higher temperature range (400-600 ºC) due to the presence of strong acid sites created by ion exchange. The strength of the acid sites increases with cerium content as the desorption peaks shift gradually towards the higher temperatures with increasing cerium in the catalysts (Barman et al., 2005). Moreover, the number of acid sites also increases as the peak area increases with Ce exchange. Thus, it can be concluded from Fig. 3 that cerium exchange leads to an increase in both quantity and strength of the acid sites in zeolite (Thakur et al., 2014; Wang and Zou, 2003). The acidity of H-beta zeolite was found to be 0.88 mmol/gm and that of cerium beta zeolite was 1.79 mmol/gm.

Figure 3 Ammonia-TPD profile of different beta zeolites. 

The FTIR spectra of H-beta and cerium exchanged beta zeolite are shown in Fig. 4. The IR spectra of H-beta and CeB10 zeolite show major bands in the 450-1650 cm−1 region with additional bands at the 3450 cm−1 region. Beta zeolite shows bands approximately at 460, 520, 785, 1060, 1212, 1625 and 3440 cm−1. The peak at 520 cm−1 is characteristic of beta zeolite (Kiri et al., 1994). The major characteristic bands at 460, 785, 1060 and 1212 cm−1 are attributed to a T-O bending mode, O-T-O external symmetric stretch and O-T-O internal asymmetric stretching vibrations and O-T-O external asymmetric stretch, respectively. In case of cerium-modified beta zeolite all the peaks show a little shift in peak positions. The additional absorption band at wavenumber 553 cm-1 represents the Ce-O stretch (Ansari et al., 2009). The intensities of peaks for the modified zeolite are greatly enhanced. The broad bands observed in the case of both catalysts at 3453 cm−1 are due to the stretching vibrations of hydroxyl groups, and those at 1635 cm−1 are lattice water molecules, respectively. This suggests that the zeolite framework remains unaffected after the ion exchange process and cerium exchange of beta zeolite.

Figure 4 FTIR of H-beta and Cerium modified beta zeolites. 

Activity of Catalyst in the Transalkylation Reaction

The activity of Ce-beta zeolites and the parent H-beta zeolite were tested for a 3.5 h time period on stream at 573 K and atmospheric pressure. Fig. 5 shows that a maximum of 83.82% selectivity of cumene was obtained over CeB10, which decreases to a value of 81.09% after 3.5 h on stream. This may be due to deactivation of zeolite by formation of coke, which blocks the catalyst surface. It is observed that the selectivity of cumene decreases with a decrease in cerium content of zeolite and was found to be 81.33% and 77.87% over CeB8 and CeB6 zeolites, respectively. Selectivity of cumene becomes lowest (73.23%) in the case of unmodified beta zeolite.

Figure 5 Cumene selectivity as a function of time-on-stream. Reaction Conditions: Pressure =1 atm, temperature = 573 K, benzene/1,4-DIPB = 5:1, space time = 9.03 kg h/ kmol, N2 to feed ratio = 0.2, catalyst amount = 0.002 kg. 

Effect of Temperature on DIPB Conversion and Product Selectivity

The effect of temperature on product selectivity is shown in Fig. 6 in the temperature range of 493 K-593 K over CeB10 catalyst. With the increase in temperature, the cumene selectivity increases gradually and reaches a maximum (83.82%) at 573 K but above this temperature the cumene selectivity is decreased due to the formation of the side product xylene.

Figure 6 Effect of temperature on product selectivity (Reaction Conditions: Pressure =1 atm, benzene/ 1,4-DIPB ratio = 5:1, space time = 9.03 kg h/ k mol, N2 to feed ratio = 0.2, time on stream = 30 min, catalyst (CeB10) amount = 0.002 kg). 

The product distribution at different reaction temperatures is shown in Table 1. The conversion of DIPB was maximum (94.69%) at a temperature of 573 K and then decreased to 85.12% at 613 K. At higher temperature the decrease in conversion is due to deactivation of the catalyst by deposition of coke on the catalyst surface.

Table 1 Product distribution at different tem­peratures. 

Product distribution (wt%) Temperature (K)
493 513 533 553 573 593
Aliphatics 0.76 0.79 0.72 0.40 0.07 0.03
Benzene 67.15 67.07 66.91 63.27 47.21 56.40
Toluene 0.12 0.20 0.22 0.30 0.35 0.69
Xylene 0.17 0.23 0.31 0.74 1.14 1.55
Cumene 2.96 10.60 15.03 20.18 42.94 32.54
n-PB 0.29 0.35 0.41 0.50 0.58 1.56
C10 1.39 2.41 5.31 5.16 0.97 0.53
1,3-DIPB 6.10 5.89 5.72 5.66 5.17 3.33
1,4-DIPB 21.06 12.46 5.40 3.79 1.57 3.37
Cumene yield (%) 2.96 10.60 15.03 20.18 42.94 32.54
1,4-DIPB conv(wt%) 29.01 51.69 81.78 87.20 94.69 88.63

Reaction Conditions: Pressure =1 atm, benzene/1,4-DIPB ratio = 5:1, space time = 9.03 kg h/ kmol, N2 to feed ratio = 0.2, time on stream = 30 min, catalyst (CeB10) amount =0.002 kg.

Effect of Benzene to 1,4-DIPB Mole Ratio on Product Selectivity

In the transalkylation reaction, the benzene to 1,4 DIPB ratio was varied from 1 to 15 at a reaction temperature of 573 K and space time of 9.03 kg h/k mol. Fig. 7 shows that the maximum selectivity (83.82%) of cumene was obtained at a benzene to 1,4-DIPB ratio of 5:1, beyond this ratio, the isomerisation of 1,4-DIPB seems to increase and hence a decrease in cumene selectivity was observed.

Figure 7 Effect of reactant mole ratio on product selectivity (Reaction Conditions: Pressure =1 atm, temperature = 573 K, space time = 9.03 kg h/ kmol, N2 to feed ratio = 0.2, time on stream = 30 min, catalyst (CeB10) amount = 0.002 kg). 

Effect of Space Time on Selectivity of Products and Conversion of 1,4 DIPB

The effect of space-time was studied in the range of 4.21 kg h/k mol-9.03 kg h/k mol. The cumene selectivity increased with increase in space time and reached its maximum of 83.82% at a space time of 9.03 kg h/kmol. Higher contact time allows the reactants to remain in the vicinity of each other and the catalyst surface for longer time, therefore enhancing the transalkylation reaction instead of isomerisation and disproportionation. Selectivity of products as a function of space time is shown in Figure 8.

Figure 8 Effect of space time on product selectivity (Reaction Conditions: Pressure =1 atm, temperature = 573 K, benzene/1,4-DIPB ratio = 5:1, N2 to feed ratio = 0.2, time on stream = 30 min, catalyst (CeB10) amount = 0.002 kg). 

The product distribution at various space times is given in Table 2. As can be seen from Fig. 9, the conversion of 1,4-DIPB increased with an increase in space time; which may be due to the higher contact time between reactants and catalyst. Conversion of 1,4-DIPB was increased to 94.69% at a space time of 9.03 kg h/kmol.

Table 2 Product distribution at various space times. 

Product distribution (wt%) Space Time (kg h/ k mol)
4.2 4.8 5.75 7.02 7.9 9.03
Aliphatics 0.71 0.49 0.48 0.04 0.10 0.06
Benzene 68.82 68.25 67.91 62.42 53.97 47.26
Toluene 0.94 0.58 0.29 0.48 0.22 0.35
Xylene 0.32 0.39 0.42 1.03 0.56 1.14
Cumene 16.96 18.78 21.02 28.65 37.39 42.93
n-PB 0.11 0.16 0.23 0.36 0.52 0.57
C10 3.34 2.42 1.14 0.14 0.70 0.96
1,3-DIPB 2.08 4.01 4.29 4.49 4.66 5.16
1,4-DIPB 6.72 4.92 4.22 2.39 1.88 1.57
Cumene yield (%) 16.96 18.78 21.02 28.65 37.39 42.93
1,4-DIPB conv.(wt%) 77.34 83.41 85.77 92.39 93.69 94.69

Reaction Conditions: Pressure =1 atm, temperature = 573 K, benzene/ 1,4-DIPB ratio = 5:1, N2 to feed ratio = 0.2, time on stream = 30 min, catalyst (CeB10) amount = 0.002 kg.

Figure 9 Effect of space time on 1,4-DIPB conversion at different temperatures. Reaction Conditions: Pressure = 1 atm, benzene/1,4-DIPB ratio = 5:1, N2 to feed ratio = 0.2, time on stream = 30 min, catalyst (CeB10) amount = 0.002 kg. 


In solid catalyzed gas phase reactions, the effect of external and internal mass transfer resistances may be significant. Experiments were carried out to find the importance of external diffusional resistance by varying feed rates and catalyst size at constant space-time. Table 3 shows that the conversions of 1,4-DIPB at constant space time are independent of feed rate. Therefore, the external mass transfer resistance is negligible. To investigate the effect of intraparticle diffusion the experiments were carried out keeping space-time constant but varying the catalyst particle size. The experimental data shown in Table 4 indicate that the conversion of 1,4-DIPB remain same with catalyst size which indicates negligible intraparticle mass transfer resistance in the particle size range studied.

Table 3 Effect of external diffusional resistances on conversion of 1,4-DIPB over CeB10

Space-time (kg h/kmol) Conversion of 1,4-DIPB (%)
Catalyst weight = 0.002 kg Catalyst weight =0.004 kg
4.2 64.34 65.11
4.8 69.80 70.87
5.75 74.03 75.12

Reaction Conditions: Pressure = 1 atm, temperature = 573 K, benzene/ 1,4-DIPB ratio = 5:1, N2 to feed ratio = 0.2, time on stream = 30 min.

Table 4 Effect of intraparticle diffusional resistance on DIPB conversion over CeB10. 

Particle size dp (mm) Conversion of 1,4-DIPB (%)
space-time (kg h/kmol) =4.2 space-time (kg h/kmol) = 4.8 space-time (kg h/kmol)= 5.75
0.50 65.87 70.23 75.34
1.00 65.35 69.98 74.76
1.50 64.34 69.80 74.03

Reaction Conditions: Pressure =1 atm, temperature = 573 K, benzene/ 1,4-DIPB ratio = 5:1, N2 to feed ratio = 0.2, time on stream = 30 min.


The kinetic runs were carried out at four different temperatures. The experiments were conducted in the zone in which mass transfer effects were negligible. Based on the product distribution various kinetic models (adsorption, desorption, and surface reaction control) were formulated following the Langmuir-Hinshelwood approach. These models were tested with the help of the experimental data. A non-linear regression algorithm was used for the kinetic parameter estimation. All models, except one, surface reaction control, were rejected as they resulted in negative kinetic constants. The following rate equation based on surface reaction-control was found to fit the experimental data satisfactorily. Transalkylation of 1,4 DIPB with benzene is a complex reaction which is followed by isomerization and disproportionation reactions.

  1. 1,4-DIPB transalkylation:


  2. Isomerisation:


  3. Dispropotionation:


  4. (4)

For the above reactions, the possible rate equations based on different mechanisms are presented below. k4 is not considered while developing the model because the model is in terms of conversion of 1,4-DIPB and this reaction does not involves DIPB. k3 is also not considered since only those reactions whose product yield is significant are taken into considerations.

Dual-site mechanism:




Single-site mechanism:




Stoichiometric model:


The partial pressure of 1,4 DIPB and benzene are related to the fractional conversions and the total pressure (P) by these following equations:




The optimum values of the parameters were obtained by minimizing the objective function given by the equation:


Model Selection

By using the values of the constants for Equation (5) for the dual site mechanism, as shown in Table 5, the standard error of estimate for the rate of disappearance of 1,4-DIPB was ±3.14 x 10-4. For Equation (6), with the values of the constants from Table 6, the standard error was ± 2.41 x 10-3. For Equation (7), with the values of the constants from Table 7, the standard error was ± 9.92 x 10-3. By comparing the standard errors, model Equation (5) was considered to be the best for representing the reaction system under investigation. The experimental and the predicted 1,4 DIPB conversions from Equation (5) at four different temperatures are plotted in Fig. 10. The figure shows that the proposed reaction rate expression predicts 1,4-DIPB conversion values quite comparable to the experimental values, having an R2 value of 0.99. The kinetic constants evaluated and tabulated at various temperatures (Table 5) were used to determine the activation energy for cumene synthesis (transalkylation) which was estimated to be 116.53 kJ/mol; for the isomerisation reaction the activation energy was found to be 176.01 kJ/mol by the following equation:


Table 5 Kinetic and adsorption constants at dif­ferent temperature. 

Kinetic and adsorption parameters Temperature (K)
493 513 533 553
k1 (k mol/kg h) 1.11 2.34 10.49 20.53
k2 (k mol/kg h) 0.097 0.534 4.54 10.54
K Dipb (atm-1) 1.223 0.839 0.689 0.459
K B (atm-1) 2.345 2.00 1.80 1.50

Figure 10 Experimental vs predicted rate of reaction. 

The plot of ln k vs 1/T is represented in Fig. 11. The activation energy values for various reactions compare well with the values of similar reactions on zeolites obtained by other investigators (Kondamudi and Upadhyayula, 2008). The kinetic model developed was able to predict the product distribution. Formation of cumene and 1,3-DIPB, the main products, is derived from the equations below.

Figure 11 A plot of ln k vs 1/T for determination of the activation energy. 

Rate of formation of cumene:


Rate of formation of 1,3-DIPB:


k1, k2, kB, kDIPB were estimated by using nonlinear regression analysis using the least square criterion. pB and pDIPB values can be calculated by the Equations (10) and (11). By using the known parameters k1, k2, kB, kDIPB, pB, and pDIPB, in Equations (15) and (16), formation of cumene and, in Equation (2), formation of 1,3-DIPB can be predicted.


The catalytic performance of cerium exchanged beta zeolite was evaluated for the commercially important cumene synthesis. H-beta zeolite was successfully modified with cerium and it was found that activity and selectivity of the cerium-modified zeolite was higher than the parent zeolite due to the presence of strong acid sites in the exchanged zeolite. 94.69% conversion of 1,4-DIPB was obtained over CeB10 zeolite with 8.3 wt% Ce content, at 573 K and a benzene to 1,4-DIPB reactant ratio of 5:1. The cumene selectivity in this condition was 83.82%. Based on the product distribution patterns, a kinetic model was proposed and the parameters of the model were estimated. The activation energies for the transalkylation and isomerisation reactions were found to be 116.53 kJ/mol and 176.01 kJ/mol respectively. In industry, using this catalyst, the rate of reaction can be increased, which can save the fixed cost as well as the operating cost by minimizing energy consumption, reactor volume, and reaction time and make the process more economic. The kinetic model derived for the reactions can provide the necessary information to scale up the reactor for large scale production in industry.


B benzene
C cumene
DIPB diisopropyl benzene
k1, k2, k3, k 4 kinetic constant (kmol/kg atm2 h)
KB adsorption constant for benzene (atm-1)
KDIPB adsorption constant for DIPB (atm-1)
KC adsorption constant for cumene (atm-1)
P total pressure (atm)
PDIPB partial pressure of DIPB (atm)
pB partial pressure of benzene (atm)
pC partial pressure of cumene (atm)
τ space-time (kg h/kmol)
XDIPB fractional conversion of DIPB
XB moles of benzene reacted (kmol)
XC moles of cumene formed (kmol)
Xexpt experimentally measured fractional conversion of DIPB
Xpred model predicted fractional conversion of DIPB


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Received: May 26, 2015; Revised: September 25, 2015; Accepted: October 29, 2015

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