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HZSM-5 catalyst
Fuel Processing Technology 176 (2018) 197–204
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Using two-zone fluidized bed reactor in propane aromatization over Zn/ HZSM-5 catalyst Abbas Roshanaei, Seyed Mehdi Alavi
T
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Reaction Engineering Lab, Chemical Engineering Department, Iran University of Science and Technology, Tehran, Iran
A R T I C LE I N FO
A B S T R A C T
Keywords: Propane aromatization Zn/HZSM-5 zeolite Catalyst deactivation Two-zone fluidized bed reactor
Propane conversion to aromatics was carried out in a two-zone fluidized bed reactor (TZFBR) over Zn/HZSM-5 catalyst. At the steady state condition the coke formation in the TZFBR was counteracted with continuous catalyst regeneration by oxygen. In the lower part of the reactor, the oxidation and combustion of deposited coke on the catalyst surface were achieved with a necessary amount of oxygen. The aromatization reaction takes place in the upper part. In order to ensure the stability of the TZFBR performance, the propane aromatization was carried out in the conventional fluidized and fixed bed reactors in the absence of oxidant. Because of the catalyst regeneration, the high propane conversion and selectivity to aromatics were obtained in the TZFBR. The effects of the main operating variables such as: O2 mole percentage, height of the propane feed entry point, bed temperature and relative velocity were studied and the optimum experimental conditions were considered.
1. Introduction Transformation of lower cost liquefied petroleum gas (LPG), mainly composed of propane and butane, to more valuable products is an important area in the petrochemical industry. In this regard, catalytic conversion of light hydrocarbons especially propane and butane to aromatics has received considerable attention from both academic and industrial point of view. BTX aromatic hydrocarbons, namely, benzene, toluene and xylenes are the most important building blocks for the petrochemicals and polymers [1–5]. Many studies have reported that the ZSM-5 zeolite modified by metals such as zinc [6–8] or gallium [9–11] is considered as an effective catalyst for the aromatization of light hydrocarbons. During the conversion of light alkanes into aromatics over ZSM-5 zeolite with metals, different types of reactions occur: cracking, dehydrogenation, oligomerization, cyclization, aromatization and coke formation [1,12–14]. Coke formation reactions are undesirable which leads to catalyst deactivation. Carbonaceous deposits block the ZSM-5 zeolite pores and the access of reactant to the active sites of zeolite being restricted. However, the deactivated zeolite can be regenerated by oxidant under controlled conditions [15–17]. Several industrial processes have been developed for conversion of light alkanes (propane and butanes) to the aromatics, including M-2 forming from Mobil, Cyclar from BP and UOP jointly, Aroforming from IFP and SALUTEC, and Z-Forming from Mitsubishi. The ZSM-5 zeolite with and without an activating agent was utilized as the chosen catalyst
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Corresponding author. E-mail address: [email protected] (S.M. Alavi).
https://doi.org/10.1016/j.fuproc.2018.03.020 Received 28 January 2018; Received in revised form 3 March 2018; Accepted 20 March 2018 0378-3820/ © 2018 Elsevier B.V. All rights reserved.
in all of aforementioned processes. These processes require catalyst regeneration, due to coke formation and the subsequent catalyst deactivation. In the M-2 forming, Aroforming and Z-Forming processes, cyclic regeneration (CR) technology and in the Cyclar process, continuous catalyst regeneration (CCR) technology are used for catalyst regeneration [2,8,18–21]. To remove the coke deposits from the catalyst surface in the light hydrocarbon aromatization process, CCR type or swing type reactor unit construction are required, which is associated with a significant economic burden. Two-zone fluidized bed reactor (TZFBR) is a multifunctional reactor in which the main reaction in reaction zone as well as the coke deposits on the catalyst surface oxidation in regenerative zone can occur simultaneously in a single fluidized bed. Hydrocarbons through intermediate part and the oxidant with amount of inert gas from the bottom of the reactor enter the TZFBR. Thus, the reaction zone at the upper part and the catalyst regeneration zone at the lower part of the reactor are created [22–24]. The excellent solid mixing in fluidized beds ensures that the regenerated catalyst at the bottom part and the deactivated catalyst at the top part can be well exchanged between two zones of TZFBR. Moreover, the produced heat by the combustion of coke in the regeneration zone is transported to the reaction zone by solids, in which the endothermic aromatization reaction occurs. This provides a good isothermicity in TZFBR and avoids hot spots. In recent years, TZFBR has been studied for the various reactions: steam reforming of methane [25], oxidative dehydrogenation of n-butane [26] and butene [27], the catalytic
Fuel Processing Technology 176 (2018) 197–204
A. Roshanaei, S.M. Alavi
Nomenclature Ar dp DTA GHSV H hh ni T
TG TZFBR U Umf Ur XO2 μg ρg ρs
Archimedes number average particle diameter, cm differential thermal analysis gas hourly space velocity, cm3 g−1 h−1 height of the bed, cm height of the propane feed entry point, cm number of carbon atoms for hydrocarbon i bed temperature, °C
thermo gravimetric two-zone fluidized bed reactor gas velocity, cm/s Minimum fluidization velocity, cm/s relative velocity oxygen mole percentage in the feed, % gas viscosity, g cm−1 s−1 gas density, g cm−3 solid density, g cm−3
2. Experimental
dehydrogenation of propane and n-butane [23,28], the n-butane partial oxidation to maleic anhydride [29] and methane aromatization [22]. Several studies regarding reactor size [30], bubble formation [31,32] and TZFBR hydrodynamics [32,33] have been performed. Based on aforementioned, it seems that TZFBR is a suitable reactor for propane aromatization over ZSM-5 zeolite at steady state operation. However to our knowledge, a two-zone fluidized bed reactor has not been investigated for the propane aromatization. This study investigates the propane aromatization over Zn/HZSM-5 catalyst in a TZFBR. At first, the stability of the TZFBR system was investigated. Then, an experimental design was developed to investigate the effects of temperature, relative velocity, mole percentage of O2 in the feed and their interactions on the TZFBR performance. In addition, the effects of the main variables in the TZFBR such as: mole percentage of O2 in the feed, height of the propane feed entry point, relative velocity and temperature were studied individually.
2.1. Catalyst synthesis ZSM-5 zeolite was prepared by hydrothermal method. Na-ZSM-5 was obtained from Ludox (30 wt% SiO2), sodium aluminate, sodium hydroxide and deionized water in a stainless steel autoclave at 170 °C for 144 h. The resulting synthesized sample was calcined at 550 °C for 16 h. NH4-ZSM-5 was obtained by Na-ZSM-5 ion exchange with ammonium nitrate (1 M) solution twice under reflux at 70 °C for 24 h. Then, the synthesized sample was calcined in air at 600 °C for 5 h to provide HZSM-5 (Si/Al ratio = 26). Zinc was introduced on HZSM-5 by wet impregnation method. 3wt.%Zn/HZSM-5 catalyst was prepared by wet impregnation of zinc nitrate solution on the HZSM-5. The sample was dried and calcined at 600 °C in air for 5 h. A detailed description of catalyst preparation was provided in the previous article [34].
Fig. 1. Schematic diagram of the experimental setup. 198
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0.8 cm/s was determined. In addition, the minimum fluidization velocity can be calculated via theoretical equations, e.g. Grace correlation [35], Eq. (1):
2.2. Catalyst characterization The crystalline structure of the synthesized catalyst was characterized by X-ray diffraction (XRD) measurements, before and after the reaction. The XRD patterns were collected using a PANalytical diffractometer with CuKα radiation (λ = 1.5406 Å, 40 kV, 40 mA) and the 2θ ° scan range between 5° to 80° with a 2θ-step of 0.026°. The crystal size and morphology of the prepared ZSM-5 catalyst were studied using scanning electron microscopy (SEM) model VEGA TESCAN. Prior to undertaking SEM image, the sample was coated with gold. The spent catalyst after 40 h time on stream was thermally analyzed to study the two-zone fluidized bed reactor performance accuracy and separation of reaction and regeneration zones in the reactor. Thermo gravimetric (TG) and differential thermal analysis (DTA) using a BAHR, STA504 (Germany) were carried out under static air with heating rate of 10 °C/min, starting from 50 °C to 800 °C.
dp Umf ρg μg
27.22 + 0.0408Ar − 27.2
=
(1)
where Ar =
dp3ρg (ρs − ρg ) g μg 2
is Archimedes number, ρg is gas density, ρs is
solid density, μg is gas viscosity and dp is the average particle diameter. Using the Grace formula, the calculated Umf in the case of Zn/HZSM-5 catalyst is about 0.5 cm/s which this result is 37% less than the Umf determined experimentally. The main reactor related variables included mole percentage of O2 in the feed (XO2), height of the propane feed entry point (hh), relative velocity (Ur = U/Umf) and bed temperature (T). The variables range with the reference values used in this study is shown in Table 1. In all runs, 55 g of catalyst and 50 mol% of propane in feed were used. Gas hourly space velocity (GHSV) is defined the total volume flow rate of feed over the mass of catalyst. By shifting relative velocity from 1.3 to 2.0, the gas hourly space velocity (GHSV) was changed from about 1350 to 2050 cm3 g−1 h−1, respectively. It is remarkable that the considered mole percentage values of TZFBR feed, throughout this research, were those that would be achieved when feed was premixed. Also, the experiments were carried out under reference conditions given in table1. After performing the TZFBR experiments, the catalyst bed enabled to cool under an inert nitrogen atmosphere. Propane conversion, product selectivity and yield are defined as follows:
2.3. Reaction system A scheme of the experimental system employed for propane aromatization process is shown in Fig. 1. The fluidized bed reactor was a stainless steel tube 25 cm long and inner diameter of 3.5 cm which was equipped with a gas distributor. The gas distributor involves a fixed bed of sands which sandwiched between two layers of stainless steel perforated plates. A section with a diameter of 5 cm and 10 cm long was placed at top part of reactor to settle the entrained catalyst particles. A stainless steel tube with an inner diameter of 3 mm served as axially movable propane inlet and was placed to introduce the hydrocarbon feed at different reactor height. Additionally an axially movable K-type thermocouple was used to measuring temperature at different parts of the reactor. An O2-N2 mixture was fed in the bottom of the reactor by gas distributor after passing from pre-heater section (inner diameter (i.d.) = 3.5 cm, long = 75 cm) and reaching to 300 °C. An external electrical heater with a PID controller was used. In all runs, a near constant temperature profile in the two-zone fluidized bed was obtained. All the streams were fed by the mass flow controllers (MFCs) to the bed. It is noteworthy that the mole percentage of O2 is in the total feed (N2 + O2 and C3H8). In addition, the propane aromatization process was carried out in the conventional fluidized bed and fixed bed reactors in the absence of oxygen and compared with two-zone fluidized bed reactor at the same operation conditions to ensure the performance stability of TZFBR. A propane‑nitrogen mixture (50 mol% propane) was fed in the bottom of the conventional fluidized bed reactor. The fixed bed reactor consisted of a quartz tube (inner diameter (i.d.) = 6 mm, outer diameter (o.d.) = 8 mm) installed in a vertical furnace. A mixture of propane and nitrogen (50 mol% propane) was introduced by two mass flow controllers from top of the fixed bed reactor. 300 mg of the catalyst (particle fraction between 0.3 and 0.6 mm) was loaded in the fixed bed reactor. The prepared catalyst was packed between two layers of quartz wool in order to avoid catalyst particles movement. The reaction products were analyzed by an on-line Thermo Finnigan (Model No. KAV00109) gas chromatograph (GC), equipped with Hayesep Q, MolSieve 13× columns and thermal conductivity detector (TCD) for detection of permanent gases such as nitrogen, carbon oxides and Q-plot and flame ionization detector (FID) for separation of the hydrocarbon products. The outlet pipeline from the reactor to the GC was heated at about 160 °C to avoid condensation of aromatics and heavy products. The zeolite catalyst was sieved to a mean particle diameter of 116 μm. Minimum fluidization velocity, Umf, was estimated with N2 at 500 °C by the bed pressure drop curve versus gas velocity. Umf was obtained from the intercept of the line of pressure drop versus decreasing gas velocity (backward) with the maximum pressure drop. According to the results of Fig. 2, the minimum fluidization velocity of
Propane conversion =
Selectivity of i =
Yield of i =
moles of propane consumed × 100, % moles of propane introduced
ni × moles of i formed × 100, % 3×moles of propane consumed
ni × moles of i formed × 100, % 3×moles of propane introduced
(2)
(3)
(4)
where ni is the number of carbon atoms for compound i in hydrocarbon products. 3. Results and discussion 3.1. Catalyst characterization The XRD patterns of the fresh and used catalyst under different operation conditions are shown in Fig. 3. Fig. 3 shows that all the characteristic peaks of fresh catalyst exist in the XRD of spent catalyst which indicates that the crystalline structure of catalyst was maintained after the reaction. Moreover the XRD patterns indicate that both Zn (2θ = 34.4–37.2°) and ZSM-5 phases (2θ = 7.5–9° and 2θ = 22.5–24.5°) are present in the catalyst before and after the reaction.
Fig. 2. Determination of Umf for Zn/HZSM-5 catalyst. 199
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reactor than in the fixed bed reactor. This result suggests that more coke was formed in the fluidized bed reactor at the same operation conditions. The results show that the highest propane conversion and also selectivity to aromatics were attained in two-zone fluidized bed reactor than the fluidized and fixed bed reactors because of the catalyst regeneration.
Table 1 Range of the operating conditions used. Variable
Reference value
Range studied
Particle diameter, dp Solid density, ρs Temperature, T Relative velocity, Ur Oxygen mole percentage, XO2 Height of the propane feed entry, hh Height of the bed, H Catalyst weight
116 μm 1.8 g/cm3 530 °C 1.6 1.5% 5 cm 10 cm 55 g
– – 500–560 °C 1.3–2 1–3% 3.5–6.5 cm – –
3.3. Analysis of the experimental two-level factorial design A two-level factorial design [37] was applied to investigate the effects of operating variables such as: temperature, O2 mole percentage in feed, relative velocity and their interactions on the performance of the TZFBR in the propane aromatization. The higher and lower values of each factor are listed in Table 2 which were chosen based on the literature data and experimental set-up limitations. The obtained amounts of the propane conversion, product selectivity and yield of aromatics under different reaction conditions based on the factorial design are shown in Table 3. The analysis of variance (ANOVA) was used to determine the effects of the factors and interactions between different factors. The results of ANOVA analysis for propane conversion and selectivity to aromatics as responses are given in Table 4. All the experiments were carried out two times and until 4% deviations for responses observed. In order to determine the significance of each parameter, both individually and interacting, the Fischer test (F-test) was used [37,38]. In the F-test, the F-value will be compared F-critical values. If F-value is larger than the critical values (F = 5.3 at 95% confidence level), then the corresponding effect is statistically significant [39]. The results of Table 4 confirm that the temperature factor is highly significant for TZFBR performance. The interaction between the temperature and relative velocity on the propane conversion and the interaction between temperature and O2 mole percentage on the selectivity to aromatics are significant, while the other interactions are not significant. In addition, temperature, O2 mole percentage and relative velocity factors are significant (Table 4). Hence, the effects of variables on TZFBR performance will be considered in next sections.
Fig. 3. XRD patterns of (a) fresh Zn/HZSM-5 catalyst (b) used catalyst in the two-zone fluidized bed reactor after several hours on reaction under different operation conditions.
The SEM image of the synthesized catalyst is shown in Fig. 4. Fig. 4 shows the cubic crystallites with a size of about 1–6 μm. Such a ZSM-5 zeolite catalyst morphology has been reported [36]. TG and DTA curves of the spent catalyst after 40 h time on stream under different operation conditions are shown in Fig. 5. The catalyst weight loss with exothermic DTA peak due to oxidation of coke at high temperature is not observed. It indicates the absence of coke deposited on the spent catalyst and shows that catalyst regeneration in two-zone fluidized bed reactor was performed correctly.
3.4. Effects of the oxygen mole percentage in feed 3.2. Study of the stability of the TZFBR system
Fig. 8 shows the effect of mole percentage of O2 in the feed on the propane conversion, yield of aromatics, selectivity to COx and aromatics in the TZFBR performance. A constant input of propane was applied in these experiments, while the amount of oxygen and nitrogen were differed. The results of Fig. 8 show that the propane conversion and COx selectivity increased with the oxygen mole percentage. Since oxygen was spent for more coke burning, higher catalyst activity
Fig. 6 shows the evolution of propane conversion, selectivity to COx and aromatics vs. time on stream for reference operating conditions (given in Table 1). Before the reaction, the catalyst was heated by 10 °C/min in a flow of N2 + O2 up to reaction temperature. Then feed propane was introduced to the TZFBR and the catalyst was kept at reaction temperature for 30 min. Unsteady state behavior was observed for initial hours. At the initial hours propane conversion and selectivity to aromatics decreased due to the coke formation on the catalyst and consequently COx selectivity was increased which is appeared as a product of coke oxidation in the regeneration zone. It can be seen that after approximately 5 h, the propane conversion and product selectivity are constant, which means that a dynamic equilibrium is achieved between coke formation in the reaction zone and coke burning in the regeneration zone. Also cracking and oligomerization products such as olefins: ethylene, propylene and butene, paraffins: methane, ethane and butane, heavy products: olefins and paraffins with 5 or more carbon were detected. After that, all the results reported correspond to the steady state attained after 5 h on reaction. The evolution of propane conversion and selectivity to aromatics as a function of time for the TZFBR, fluidized and fixed bed reactors is shown in Fig. 7. Fig. 7 shows the excellent stability of the TZFBR after several hours due to regeneration of the catalyst by oxygen. The propane conversion as a consequence of catalyst deactivation via coke deposition decreased in the fluidized and fixed bed reactors with time on stream. The propane conversion decreased faster in the fluidized bed
Fig. 4. The SEM image of the synthesized Zn/HZSM-5 catalyst. 200
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Fig. 5. TG and DTA curves of the spent Zn/HZSM-5 catalyst after 40 h on reaction in the two-zone fluidized bed reactor under different operation conditions.
Fig. 7. Evolution of (a) propane conversion and (b) selectivity to aromatics with time for TZFBR, fluidized and fixed bed reactors. Operating conditions: reference values given in Table 1.
Fig. 6. Evolution of propane conversion, selectivity to COx and aromatics with time for the TZFBR at reference conditions given in Table 1.
results, selectivity and yield to aromatics enhance with increasing oxygen mole percentage from 1% to 1.5%. It can be seen that when oxygen mole percentage is higher than 1.5%, selectivity and yield to aromatics decrease whereas propane conversion and COx selectivity increase. This means that the excess amount of oxygen for coke burning entered into the reaction zone and was consumed for propane combustion to produce COx and decrease selectivity to aromatics. These results are in agreement with the results reported for propane dehydrogenation [23]. According to the obtained experimental results, the amount of oxygen fed used in the TZFBR is important for the propane aromatization process. It must be consumed for oxidizing coke formed in the propane aromatization process, and at the same time avoiding propane and oxygen mixing, because it leads to decrease selectivity to aromatics.
Table 2 Factors and levels used in the factorial design. Factors
Low level (−1)
High level (+1)
Temperature (°C) Oxygen mole percentage (%) Relative velocity
500 1.5 1.3
530 2.5 1.6
(hh/H) up to 0.65 were not applied in order to avoid reducing the reaction zone volume and selectivity to aromatics. Also, the catalyst regeneration could not be performed completely at the height ratios (hh/ H) lower than 0.35, due to more coke formation in the larger reaction zone and smaller regeneration zone. Therefore, the experiments were followed with the range of the height ratios (hh/H) between 0.35 and 0.65. Since more coke was produced in the larger reaction zone, therefore these experiments were carried out with constant feed flow of propane, while the oxygen mole percentage was varied between around 1.5% and 2.5%. These experiments were carried out in the reference operating conditions, as indicated in Table 1 except oxygen mole
3.5. Effect of the height of the propane feed entry point Fig. 9 shows the propane conversion, selectivity to COx and aromatics for the different propane feed entry points with the same total height of the bed. The height ratios of the propane feed entry point/bed 201
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Table 3 Propane conversion, product selectivity and yield of aromatics in the TZFBR. Weight of catalyst = 55 g, 50 mol% of propane in feed. No.
T(°C)
XO2 (%)
Ur
Propane conversion (%)
Sel. to aromatics (%)
CO2 Sel. (%)
CO Sel. (%)
Sel. to paraffins (%)
Sel. to heavy products (%)
Sel. to olefins (%)
Yield of aromatics (%)
1 2 3 4 5 6 7 8
500 530 530 530 500 500 500 530
1.5 2.5 2.5 1.5 2.5 2.5 1.5 1.5
1.6 1.3 1.6 1.6 1.3 1.6 1.3 1.3
38.5 54.5 51.6 47.5 43.9 43.2 39.5 50.3
16.5 36.2 29.8 37.5 18.5 13.9 22.8 41.3
4.2 8.2 9.0 5.6 8.8 9.2 5.2 4.0
1.2 2.5 2.8 1.8 3.0 3.4 1.6 1.0
42.8 32.0 37.8 33.6 38.0 39.1 39.5 33.1
1.9 2.0 3.0 2.5 2.6 2.7 2.6 2.1
33.4 19.1 17.6 19.0 29.1 31.7 28.3 18.5
6.4 19.7 15.4 17.8 8.1 6.0 9.0 20.8
Table 4 ANOVA for propane conversion and selectivity to aromatics to determine the significance of the variables in the TZFBR. Source of variation
Effect
Propane conversion
Significance
Effect
F-value
Selectivity to aromatics
Significance
F-value
Main effect T XO2 Ur
9.70 4.35 −1.85
850.53 171.05 30.94
S S S
18.28 −4.93 −5.28
1106.34 80.35 92.18
S S S
Interactions T ∗ XO2 T ∗ Ur XO2 ∗ Ur T ∗ XO2 ∗ Ur
−0.20 −1.00 0.05 −0.10
0.36 9.04 0.02 0.09
NS S NS NS
−1.48 0.18 −0.23 −1.08
7.21 0.10 0.17 3.83
S NS NS NS
S: Significant, NS: Not significant.
Fig. 8. Propane conversion, yield of aromatics, selectivity to COx and aromatics for the TZFBR as a function of oxygen mole percentage at reference values, see Table 1.
Fig. 9. Effect of height of the propane feed entry point upon propane conversion, selectivity to COx and aromatics in TZFBR at reference operating conditions.
percentage and a steady state was obtained for all the experiments. It can be observed from Fig. 9 that as the height ratio increases, the propane conversion and COx selectivity decrease. Since at lower height ratio more catalyst is in contact with propane, therefore more coke is produced which results in increasing propane conversion and COx selectivity. In addition, the results show that selectivity to aromatics achieves a maximum at hh/H = 0.5. These results could be explained with the aid of the previous kinetic study [34]. With decreasing height ratio below than 0.5 (then large reaction zone) more coke is formed on the catalyst by means of the aromatics which leads to a decreased selectivity to aromatics and to an increased COx selectivity, while by increasing height ratio above 0.5 less aromatics and coke were produced, because of lower catalyst contact with propane feed.
3.6. Effect of the relative velocity Propane conversion, selectivity to aromatics and COx as a function of relative velocity are shown in Fig. 10. Other variables are at their reference values. It can be seen that the propane conversion and selectivity to aromatics decrease with the increase in the relative velocity, because the gas contact time in the reaction zone and possibility of contact with the catalyst particles decreased. Also, increasing of the gas superficial velocity leads to increase in the bubble size which results in an increase in the mass transfer resistance between the gas in the bubbles and the catalytic emulsion phase. Increase in COx selectivity with rising of the relative velocity is shown in Fig. 10. The results reveal that an increase of the relative
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4. Conclusion For the first time, the propane aromatization over Zn/HZSM-5 catalyst in a two-zone fluidized bed reactor was accomplished. In the reactor used in this study, the reaction and regeneration zones were created by separating the oxygen and propane feeds. The results obtained in this work show that, by using a TZFBR, the catalyst deactivation as a result of coke deposition can be compensated by continuously regenerating the catalyst, thus obtaining a steady state operation in TZFBR. This work demonstrated that the yield of aromatics in the TZFBR is highest than in the conventional fluidized and fixed bed reactors at the steady state operation, where the coke formation leads to the catalyst deactivation in the conventional reactors in the absence of oxygen which results in decrease of the propane conversion and selectivity to aromatics with time. The effects of parameters such as: temperature, relative velocity, O2 mole percentage and their interactions were studied using two-level factorial design. Temperature was the main parameter affecting the two-zone fluidized bed reactor performance. Also, the effect of the height of propane feed entry point alongside the other aforementioned parameters was studied individually and the optimum experimental conditions were investigated. Thus, the TZFBR has several advantages over the conventional reactors regards to aromatics production: (a) Reaction and regeneration of catalyst occur in the same vessel. (b) The heat produced from coke burning in catalyst regeneration zone provides a part of the energy necessary for endothermic aromatization reaction. (c) By separating the oxygen and hydrocarbon fed the flammable mixture is avoided.
Fig. 10. Effect of relative velocity upon propane conversion, selectivity to COx and aromatics in TZFBR. Other variables are at their reference values.
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Fig. 11. Propane conversion, selectivity to COx and aromatics for the TZFBR as a function of temperature at reference operating conditions.
velocity leads to increase the gas flow rate in the bubbles, therefore the solid recirculation rate in the wakes associated with bubble increase. In order to compensate the upward movement of solid particles in the wake, a downward movement of the solids in emulsion phase has been made. Probably the downward movement of the solid particles leads to transfer more catalytic particles and deposited coke on its surface from the reaction zone to the regeneration zone which result in coke burning and increasing COx selectivity slightly.
3.7. Effect of the temperature The effect of the temperature upon the propane conversion and product selectivity is presented in Fig. 11. Since more coke was produced in the propane aromatization process at the higher temperatures, more oxygen was needed at these temperatures in order to achieve a steady-state condition in the TZFBR. Accordingly, these experiments were carried out with a constant input of propane and different amount of oxygen (1.5–3% O2 mole percentage). Based on the previous developed kinetic model for the propane aromatization process [34], because the dehydrogenation, cracking and aromatization reactions are endothermic, as the temperature was raised, the propane conversion and selectivity to aromatics increased, as could be expected. Increasing of the COx selectivity with temperature is due to burning more coke formed at the higher temperature. 203
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regeneration zones, Stud. Surf. Sci. Catal. 130 (2000) 2717–2722. [29] J. Gascón, C. Téllez, J. Herguido, M. Menéndez, Fluidized bed reactors with twozones for maleic anhydride production: different configurations and effect of scale, Ind. Eng. Chem. Res. 44 (2005) 8945–8951. [30] O. Rubio, J. Herguido, M. Menéndez, Two-zone fluidized bed reactor for simultaneous reaction and catalyst reoxidation: influence of reactor size, Appl. Catal. A:Gen. 272 (2004) 321–327. [31] I. Julián, J. Herguido, M. Menéndez, A non-parametric bubble size correlation for a Two-Section Two-Zone Fluidized Bed Reactor (TS-TZFBR), Powder Technol. 256 (2014) 146–157. [32] I. Julián, F. Gallucci, M. van Sint Annaland, J. Herguido, M. Menéndez, Hydrodynamic study of a two-section two-zone fluidized bed reactor with an immersed tube bank via PIV/DIA, Chem. Eng. Sci. 134 (2015) 238–250. [33] I. Julián, J. Herguido, M. Menéndez, Particle mixing in a two-section two-zone fluidized bed reactor. Experimental technique and counter-current back-mixing model validation, Ind. Eng. Chem. Res. 52 (2013) 13587–13596. [34] A. Roshanaei, S. Mehdi Alavi, Kinetic study of propane aromatization over Zn/ HZSM-5 zeolite under conditions of catalyst deactivation using genetic algorithm, J. Serb. Chem. Soc. (2018), http://dx.doi.org/10.2298/JSC170621007R. [35] J. Grace, Fluidized bed hydrodynamics, in: G. Hetsroni (Ed.), Chapter 8.1 in Handbook of Multiphase Systems, Hemisphere Publishing, Washington, 1982. [36] G. Roohollahi, M. Kazemeini, A. Mohammadrezaee, R. Golhosseini, Chemical kinetic modeling of i-butane and n-butane catalytic cracking reactions over HZSM-5 zeolite, AICHE J. 58 (2012) 2456–2465. [37] D.C. Montgomery, G.C. Runger, Applied Statistics and Probability for Engineers, John Wiley & Sons, 2010. [38] G.E. Box, J.S. Hunter, W.G. Hunter, Statistics for Experimenters: Design, Innovation, and Discovery, Wiley-Interscience, New York, 2005. [39] G. Mazloom, S.M. Alavi, Partial oxidation of propane over Mo1V0.3Te0.23Nb0.12Ox catalyst in a fluidized bed reactor, Part. Sci. Technol. 33 (2015) 204–212.
aromatization, Appl. Catal. A: Gen. 292 (2005) 162–170. [18] C. Song, K. Liu, D. Zhang, S. Liu, X. Li, S. Xie, L. Xu, Effect of cofeeding n-butane with methanol on aromatization performance and coke formation over a Zn loaded ZSM-5/ZSM-11 zeolite, Appl. Catal. A:Gen. 470 (2014) 15–23. [19] S. Fukase, N. Igarashi, K. Kato, T. Nomura, Y. Ishibashi, Development of light naphta aromatization process using a conventional fixed bed unit, in: M. AbsiHalabi, et al. (Ed.), Catalysts Petroleum Refining Petrochemical Industries 1995, Elsevier, Amsterdam, 1996, pp. 455–464 by. [20] M.S. Gyngazova, A.V. Kravtsov, E.D. Ivanchina, M.V. Korolenko, N.V. Chekantsev, Reactor modeling and simulation of moving-bed catalytic reforming process, Chem. Eng. J. 176 (2011) 134–143. [21] W.J.H. Dehertog, G.F. Fromen, A catalytic route for aromatics production from LPG, Appl. Catal. A: Gen. 189 (1999) 63–75. [22] M.P. Gimeno, J. Soler, J. Herguido, M. Menéndez, Counteracting catalyst deactivation in methane aromatization with a two zone fluidized bed reactor, Ind. Eng. Chem. Res. 49 (2010) 996–1000. [23] J. Gascon, C. Tellez, J. Herguido, M. Menendez, A two-zone fluidized bed reactor for catalytic propane dehydrogenation, Chem. Eng. J. 106 (2005) 91–96. [24] L. Pérez-Moreno, J. Soler, J. Herguido, M. Menéndez, Stable hydrogen production by methane steam reforming in a two zone fluidized bed reactor: experimental assessment, J. Power Sources 243 (2013) 233–241. [25] L. Pérez-Moreno, J. Soler, J. Herguido, M. Menéndez, Stable hydrogen production by methane steam reforming in a two-zone fluidized-bed reactor: effect of the operating variables, Int. J. Hydrog. Energy 38 (2013) 7830–7838. [26] J. Rischard, C. Antinori, L. Maier, O. Deutschmann, Oxidative dehydrogenation of n-butane to butadiene with Mo-V-MgO catalysts in a two-zone fluidized bed reactor, Appl. Catal. A: Gen. 511 (2016) 23–30. [27] J. Rischard, R. Franz, C. Antinori, O. Deutschmann, Oxidative dehydrogenation of butanes over Bi-Mo and Mo-V based catalysts in a two-zone fluidized bed reactor, AICHE J. 63 (2017) 43–50. [28] C. Callejas, J. Soler, J. Herguido, M. Menéndez, J. Santamaría, Catalytic dehydrogenation of n-butane in a fluidized bed reactor with separate coking and
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Contents lists available at ScienceDirect
Fuel Processing Technology journal homepage: www.elsevier.com/locate/fuproc
Using two-zone fluidized bed reactor in propane aromatization over Zn/ HZSM-5 catalyst Abbas Roshanaei, Seyed Mehdi Alavi
T
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Reaction Engineering Lab, Chemical Engineering Department, Iran University of Science and Technology, Tehran, Iran
A R T I C LE I N FO
A B S T R A C T
Keywords: Propane aromatization Zn/HZSM-5 zeolite Catalyst deactivation Two-zone fluidized bed reactor
Propane conversion to aromatics was carried out in a two-zone fluidized bed reactor (TZFBR) over Zn/HZSM-5 catalyst. At the steady state condition the coke formation in the TZFBR was counteracted with continuous catalyst regeneration by oxygen. In the lower part of the reactor, the oxidation and combustion of deposited coke on the catalyst surface were achieved with a necessary amount of oxygen. The aromatization reaction takes place in the upper part. In order to ensure the stability of the TZFBR performance, the propane aromatization was carried out in the conventional fluidized and fixed bed reactors in the absence of oxidant. Because of the catalyst regeneration, the high propane conversion and selectivity to aromatics were obtained in the TZFBR. The effects of the main operating variables such as: O2 mole percentage, height of the propane feed entry point, bed temperature and relative velocity were studied and the optimum experimental conditions were considered.
1. Introduction Transformation of lower cost liquefied petroleum gas (LPG), mainly composed of propane and butane, to more valuable products is an important area in the petrochemical industry. In this regard, catalytic conversion of light hydrocarbons especially propane and butane to aromatics has received considerable attention from both academic and industrial point of view. BTX aromatic hydrocarbons, namely, benzene, toluene and xylenes are the most important building blocks for the petrochemicals and polymers [1–5]. Many studies have reported that the ZSM-5 zeolite modified by metals such as zinc [6–8] or gallium [9–11] is considered as an effective catalyst for the aromatization of light hydrocarbons. During the conversion of light alkanes into aromatics over ZSM-5 zeolite with metals, different types of reactions occur: cracking, dehydrogenation, oligomerization, cyclization, aromatization and coke formation [1,12–14]. Coke formation reactions are undesirable which leads to catalyst deactivation. Carbonaceous deposits block the ZSM-5 zeolite pores and the access of reactant to the active sites of zeolite being restricted. However, the deactivated zeolite can be regenerated by oxidant under controlled conditions [15–17]. Several industrial processes have been developed for conversion of light alkanes (propane and butanes) to the aromatics, including M-2 forming from Mobil, Cyclar from BP and UOP jointly, Aroforming from IFP and SALUTEC, and Z-Forming from Mitsubishi. The ZSM-5 zeolite with and without an activating agent was utilized as the chosen catalyst
⁎
Corresponding author. E-mail address: [email protected] (S.M. Alavi).
https://doi.org/10.1016/j.fuproc.2018.03.020 Received 28 January 2018; Received in revised form 3 March 2018; Accepted 20 March 2018 0378-3820/ © 2018 Elsevier B.V. All rights reserved.
in all of aforementioned processes. These processes require catalyst regeneration, due to coke formation and the subsequent catalyst deactivation. In the M-2 forming, Aroforming and Z-Forming processes, cyclic regeneration (CR) technology and in the Cyclar process, continuous catalyst regeneration (CCR) technology are used for catalyst regeneration [2,8,18–21]. To remove the coke deposits from the catalyst surface in the light hydrocarbon aromatization process, CCR type or swing type reactor unit construction are required, which is associated with a significant economic burden. Two-zone fluidized bed reactor (TZFBR) is a multifunctional reactor in which the main reaction in reaction zone as well as the coke deposits on the catalyst surface oxidation in regenerative zone can occur simultaneously in a single fluidized bed. Hydrocarbons through intermediate part and the oxidant with amount of inert gas from the bottom of the reactor enter the TZFBR. Thus, the reaction zone at the upper part and the catalyst regeneration zone at the lower part of the reactor are created [22–24]. The excellent solid mixing in fluidized beds ensures that the regenerated catalyst at the bottom part and the deactivated catalyst at the top part can be well exchanged between two zones of TZFBR. Moreover, the produced heat by the combustion of coke in the regeneration zone is transported to the reaction zone by solids, in which the endothermic aromatization reaction occurs. This provides a good isothermicity in TZFBR and avoids hot spots. In recent years, TZFBR has been studied for the various reactions: steam reforming of methane [25], oxidative dehydrogenation of n-butane [26] and butene [27], the catalytic
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Nomenclature Ar dp DTA GHSV H hh ni T
TG TZFBR U Umf Ur XO2 μg ρg ρs
Archimedes number average particle diameter, cm differential thermal analysis gas hourly space velocity, cm3 g−1 h−1 height of the bed, cm height of the propane feed entry point, cm number of carbon atoms for hydrocarbon i bed temperature, °C
thermo gravimetric two-zone fluidized bed reactor gas velocity, cm/s Minimum fluidization velocity, cm/s relative velocity oxygen mole percentage in the feed, % gas viscosity, g cm−1 s−1 gas density, g cm−3 solid density, g cm−3
2. Experimental
dehydrogenation of propane and n-butane [23,28], the n-butane partial oxidation to maleic anhydride [29] and methane aromatization [22]. Several studies regarding reactor size [30], bubble formation [31,32] and TZFBR hydrodynamics [32,33] have been performed. Based on aforementioned, it seems that TZFBR is a suitable reactor for propane aromatization over ZSM-5 zeolite at steady state operation. However to our knowledge, a two-zone fluidized bed reactor has not been investigated for the propane aromatization. This study investigates the propane aromatization over Zn/HZSM-5 catalyst in a TZFBR. At first, the stability of the TZFBR system was investigated. Then, an experimental design was developed to investigate the effects of temperature, relative velocity, mole percentage of O2 in the feed and their interactions on the TZFBR performance. In addition, the effects of the main variables in the TZFBR such as: mole percentage of O2 in the feed, height of the propane feed entry point, relative velocity and temperature were studied individually.
2.1. Catalyst synthesis ZSM-5 zeolite was prepared by hydrothermal method. Na-ZSM-5 was obtained from Ludox (30 wt% SiO2), sodium aluminate, sodium hydroxide and deionized water in a stainless steel autoclave at 170 °C for 144 h. The resulting synthesized sample was calcined at 550 °C for 16 h. NH4-ZSM-5 was obtained by Na-ZSM-5 ion exchange with ammonium nitrate (1 M) solution twice under reflux at 70 °C for 24 h. Then, the synthesized sample was calcined in air at 600 °C for 5 h to provide HZSM-5 (Si/Al ratio = 26). Zinc was introduced on HZSM-5 by wet impregnation method. 3wt.%Zn/HZSM-5 catalyst was prepared by wet impregnation of zinc nitrate solution on the HZSM-5. The sample was dried and calcined at 600 °C in air for 5 h. A detailed description of catalyst preparation was provided in the previous article [34].
Fig. 1. Schematic diagram of the experimental setup. 198
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0.8 cm/s was determined. In addition, the minimum fluidization velocity can be calculated via theoretical equations, e.g. Grace correlation [35], Eq. (1):
2.2. Catalyst characterization The crystalline structure of the synthesized catalyst was characterized by X-ray diffraction (XRD) measurements, before and after the reaction. The XRD patterns were collected using a PANalytical diffractometer with CuKα radiation (λ = 1.5406 Å, 40 kV, 40 mA) and the 2θ ° scan range between 5° to 80° with a 2θ-step of 0.026°. The crystal size and morphology of the prepared ZSM-5 catalyst were studied using scanning electron microscopy (SEM) model VEGA TESCAN. Prior to undertaking SEM image, the sample was coated with gold. The spent catalyst after 40 h time on stream was thermally analyzed to study the two-zone fluidized bed reactor performance accuracy and separation of reaction and regeneration zones in the reactor. Thermo gravimetric (TG) and differential thermal analysis (DTA) using a BAHR, STA504 (Germany) were carried out under static air with heating rate of 10 °C/min, starting from 50 °C to 800 °C.
dp Umf ρg μg
27.22 + 0.0408Ar − 27.2
=
(1)
where Ar =
dp3ρg (ρs − ρg ) g μg 2
is Archimedes number, ρg is gas density, ρs is
solid density, μg is gas viscosity and dp is the average particle diameter. Using the Grace formula, the calculated Umf in the case of Zn/HZSM-5 catalyst is about 0.5 cm/s which this result is 37% less than the Umf determined experimentally. The main reactor related variables included mole percentage of O2 in the feed (XO2), height of the propane feed entry point (hh), relative velocity (Ur = U/Umf) and bed temperature (T). The variables range with the reference values used in this study is shown in Table 1. In all runs, 55 g of catalyst and 50 mol% of propane in feed were used. Gas hourly space velocity (GHSV) is defined the total volume flow rate of feed over the mass of catalyst. By shifting relative velocity from 1.3 to 2.0, the gas hourly space velocity (GHSV) was changed from about 1350 to 2050 cm3 g−1 h−1, respectively. It is remarkable that the considered mole percentage values of TZFBR feed, throughout this research, were those that would be achieved when feed was premixed. Also, the experiments were carried out under reference conditions given in table1. After performing the TZFBR experiments, the catalyst bed enabled to cool under an inert nitrogen atmosphere. Propane conversion, product selectivity and yield are defined as follows:
2.3. Reaction system A scheme of the experimental system employed for propane aromatization process is shown in Fig. 1. The fluidized bed reactor was a stainless steel tube 25 cm long and inner diameter of 3.5 cm which was equipped with a gas distributor. The gas distributor involves a fixed bed of sands which sandwiched between two layers of stainless steel perforated plates. A section with a diameter of 5 cm and 10 cm long was placed at top part of reactor to settle the entrained catalyst particles. A stainless steel tube with an inner diameter of 3 mm served as axially movable propane inlet and was placed to introduce the hydrocarbon feed at different reactor height. Additionally an axially movable K-type thermocouple was used to measuring temperature at different parts of the reactor. An O2-N2 mixture was fed in the bottom of the reactor by gas distributor after passing from pre-heater section (inner diameter (i.d.) = 3.5 cm, long = 75 cm) and reaching to 300 °C. An external electrical heater with a PID controller was used. In all runs, a near constant temperature profile in the two-zone fluidized bed was obtained. All the streams were fed by the mass flow controllers (MFCs) to the bed. It is noteworthy that the mole percentage of O2 is in the total feed (N2 + O2 and C3H8). In addition, the propane aromatization process was carried out in the conventional fluidized bed and fixed bed reactors in the absence of oxygen and compared with two-zone fluidized bed reactor at the same operation conditions to ensure the performance stability of TZFBR. A propane‑nitrogen mixture (50 mol% propane) was fed in the bottom of the conventional fluidized bed reactor. The fixed bed reactor consisted of a quartz tube (inner diameter (i.d.) = 6 mm, outer diameter (o.d.) = 8 mm) installed in a vertical furnace. A mixture of propane and nitrogen (50 mol% propane) was introduced by two mass flow controllers from top of the fixed bed reactor. 300 mg of the catalyst (particle fraction between 0.3 and 0.6 mm) was loaded in the fixed bed reactor. The prepared catalyst was packed between two layers of quartz wool in order to avoid catalyst particles movement. The reaction products were analyzed by an on-line Thermo Finnigan (Model No. KAV00109) gas chromatograph (GC), equipped with Hayesep Q, MolSieve 13× columns and thermal conductivity detector (TCD) for detection of permanent gases such as nitrogen, carbon oxides and Q-plot and flame ionization detector (FID) for separation of the hydrocarbon products. The outlet pipeline from the reactor to the GC was heated at about 160 °C to avoid condensation of aromatics and heavy products. The zeolite catalyst was sieved to a mean particle diameter of 116 μm. Minimum fluidization velocity, Umf, was estimated with N2 at 500 °C by the bed pressure drop curve versus gas velocity. Umf was obtained from the intercept of the line of pressure drop versus decreasing gas velocity (backward) with the maximum pressure drop. According to the results of Fig. 2, the minimum fluidization velocity of
Propane conversion =
Selectivity of i =
Yield of i =
moles of propane consumed × 100, % moles of propane introduced
ni × moles of i formed × 100, % 3×moles of propane consumed
ni × moles of i formed × 100, % 3×moles of propane introduced
(2)
(3)
(4)
where ni is the number of carbon atoms for compound i in hydrocarbon products. 3. Results and discussion 3.1. Catalyst characterization The XRD patterns of the fresh and used catalyst under different operation conditions are shown in Fig. 3. Fig. 3 shows that all the characteristic peaks of fresh catalyst exist in the XRD of spent catalyst which indicates that the crystalline structure of catalyst was maintained after the reaction. Moreover the XRD patterns indicate that both Zn (2θ = 34.4–37.2°) and ZSM-5 phases (2θ = 7.5–9° and 2θ = 22.5–24.5°) are present in the catalyst before and after the reaction.
Fig. 2. Determination of Umf for Zn/HZSM-5 catalyst. 199
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reactor than in the fixed bed reactor. This result suggests that more coke was formed in the fluidized bed reactor at the same operation conditions. The results show that the highest propane conversion and also selectivity to aromatics were attained in two-zone fluidized bed reactor than the fluidized and fixed bed reactors because of the catalyst regeneration.
Table 1 Range of the operating conditions used. Variable
Reference value
Range studied
Particle diameter, dp Solid density, ρs Temperature, T Relative velocity, Ur Oxygen mole percentage, XO2 Height of the propane feed entry, hh Height of the bed, H Catalyst weight
116 μm 1.8 g/cm3 530 °C 1.6 1.5% 5 cm 10 cm 55 g
– – 500–560 °C 1.3–2 1–3% 3.5–6.5 cm – –
3.3. Analysis of the experimental two-level factorial design A two-level factorial design [37] was applied to investigate the effects of operating variables such as: temperature, O2 mole percentage in feed, relative velocity and their interactions on the performance of the TZFBR in the propane aromatization. The higher and lower values of each factor are listed in Table 2 which were chosen based on the literature data and experimental set-up limitations. The obtained amounts of the propane conversion, product selectivity and yield of aromatics under different reaction conditions based on the factorial design are shown in Table 3. The analysis of variance (ANOVA) was used to determine the effects of the factors and interactions between different factors. The results of ANOVA analysis for propane conversion and selectivity to aromatics as responses are given in Table 4. All the experiments were carried out two times and until 4% deviations for responses observed. In order to determine the significance of each parameter, both individually and interacting, the Fischer test (F-test) was used [37,38]. In the F-test, the F-value will be compared F-critical values. If F-value is larger than the critical values (F = 5.3 at 95% confidence level), then the corresponding effect is statistically significant [39]. The results of Table 4 confirm that the temperature factor is highly significant for TZFBR performance. The interaction between the temperature and relative velocity on the propane conversion and the interaction between temperature and O2 mole percentage on the selectivity to aromatics are significant, while the other interactions are not significant. In addition, temperature, O2 mole percentage and relative velocity factors are significant (Table 4). Hence, the effects of variables on TZFBR performance will be considered in next sections.
Fig. 3. XRD patterns of (a) fresh Zn/HZSM-5 catalyst (b) used catalyst in the two-zone fluidized bed reactor after several hours on reaction under different operation conditions.
The SEM image of the synthesized catalyst is shown in Fig. 4. Fig. 4 shows the cubic crystallites with a size of about 1–6 μm. Such a ZSM-5 zeolite catalyst morphology has been reported [36]. TG and DTA curves of the spent catalyst after 40 h time on stream under different operation conditions are shown in Fig. 5. The catalyst weight loss with exothermic DTA peak due to oxidation of coke at high temperature is not observed. It indicates the absence of coke deposited on the spent catalyst and shows that catalyst regeneration in two-zone fluidized bed reactor was performed correctly.
3.4. Effects of the oxygen mole percentage in feed 3.2. Study of the stability of the TZFBR system
Fig. 8 shows the effect of mole percentage of O2 in the feed on the propane conversion, yield of aromatics, selectivity to COx and aromatics in the TZFBR performance. A constant input of propane was applied in these experiments, while the amount of oxygen and nitrogen were differed. The results of Fig. 8 show that the propane conversion and COx selectivity increased with the oxygen mole percentage. Since oxygen was spent for more coke burning, higher catalyst activity
Fig. 6 shows the evolution of propane conversion, selectivity to COx and aromatics vs. time on stream for reference operating conditions (given in Table 1). Before the reaction, the catalyst was heated by 10 °C/min in a flow of N2 + O2 up to reaction temperature. Then feed propane was introduced to the TZFBR and the catalyst was kept at reaction temperature for 30 min. Unsteady state behavior was observed for initial hours. At the initial hours propane conversion and selectivity to aromatics decreased due to the coke formation on the catalyst and consequently COx selectivity was increased which is appeared as a product of coke oxidation in the regeneration zone. It can be seen that after approximately 5 h, the propane conversion and product selectivity are constant, which means that a dynamic equilibrium is achieved between coke formation in the reaction zone and coke burning in the regeneration zone. Also cracking and oligomerization products such as olefins: ethylene, propylene and butene, paraffins: methane, ethane and butane, heavy products: olefins and paraffins with 5 or more carbon were detected. After that, all the results reported correspond to the steady state attained after 5 h on reaction. The evolution of propane conversion and selectivity to aromatics as a function of time for the TZFBR, fluidized and fixed bed reactors is shown in Fig. 7. Fig. 7 shows the excellent stability of the TZFBR after several hours due to regeneration of the catalyst by oxygen. The propane conversion as a consequence of catalyst deactivation via coke deposition decreased in the fluidized and fixed bed reactors with time on stream. The propane conversion decreased faster in the fluidized bed
Fig. 4. The SEM image of the synthesized Zn/HZSM-5 catalyst. 200
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Fig. 5. TG and DTA curves of the spent Zn/HZSM-5 catalyst after 40 h on reaction in the two-zone fluidized bed reactor under different operation conditions.
Fig. 7. Evolution of (a) propane conversion and (b) selectivity to aromatics with time for TZFBR, fluidized and fixed bed reactors. Operating conditions: reference values given in Table 1.
Fig. 6. Evolution of propane conversion, selectivity to COx and aromatics with time for the TZFBR at reference conditions given in Table 1.
results, selectivity and yield to aromatics enhance with increasing oxygen mole percentage from 1% to 1.5%. It can be seen that when oxygen mole percentage is higher than 1.5%, selectivity and yield to aromatics decrease whereas propane conversion and COx selectivity increase. This means that the excess amount of oxygen for coke burning entered into the reaction zone and was consumed for propane combustion to produce COx and decrease selectivity to aromatics. These results are in agreement with the results reported for propane dehydrogenation [23]. According to the obtained experimental results, the amount of oxygen fed used in the TZFBR is important for the propane aromatization process. It must be consumed for oxidizing coke formed in the propane aromatization process, and at the same time avoiding propane and oxygen mixing, because it leads to decrease selectivity to aromatics.
Table 2 Factors and levels used in the factorial design. Factors
Low level (−1)
High level (+1)
Temperature (°C) Oxygen mole percentage (%) Relative velocity
500 1.5 1.3
530 2.5 1.6
(hh/H) up to 0.65 were not applied in order to avoid reducing the reaction zone volume and selectivity to aromatics. Also, the catalyst regeneration could not be performed completely at the height ratios (hh/ H) lower than 0.35, due to more coke formation in the larger reaction zone and smaller regeneration zone. Therefore, the experiments were followed with the range of the height ratios (hh/H) between 0.35 and 0.65. Since more coke was produced in the larger reaction zone, therefore these experiments were carried out with constant feed flow of propane, while the oxygen mole percentage was varied between around 1.5% and 2.5%. These experiments were carried out in the reference operating conditions, as indicated in Table 1 except oxygen mole
3.5. Effect of the height of the propane feed entry point Fig. 9 shows the propane conversion, selectivity to COx and aromatics for the different propane feed entry points with the same total height of the bed. The height ratios of the propane feed entry point/bed 201
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Table 3 Propane conversion, product selectivity and yield of aromatics in the TZFBR. Weight of catalyst = 55 g, 50 mol% of propane in feed. No.
T(°C)
XO2 (%)
Ur
Propane conversion (%)
Sel. to aromatics (%)
CO2 Sel. (%)
CO Sel. (%)
Sel. to paraffins (%)
Sel. to heavy products (%)
Sel. to olefins (%)
Yield of aromatics (%)
1 2 3 4 5 6 7 8
500 530 530 530 500 500 500 530
1.5 2.5 2.5 1.5 2.5 2.5 1.5 1.5
1.6 1.3 1.6 1.6 1.3 1.6 1.3 1.3
38.5 54.5 51.6 47.5 43.9 43.2 39.5 50.3
16.5 36.2 29.8 37.5 18.5 13.9 22.8 41.3
4.2 8.2 9.0 5.6 8.8 9.2 5.2 4.0
1.2 2.5 2.8 1.8 3.0 3.4 1.6 1.0
42.8 32.0 37.8 33.6 38.0 39.1 39.5 33.1
1.9 2.0 3.0 2.5 2.6 2.7 2.6 2.1
33.4 19.1 17.6 19.0 29.1 31.7 28.3 18.5
6.4 19.7 15.4 17.8 8.1 6.0 9.0 20.8
Table 4 ANOVA for propane conversion and selectivity to aromatics to determine the significance of the variables in the TZFBR. Source of variation
Effect
Propane conversion
Significance
Effect
F-value
Selectivity to aromatics
Significance
F-value
Main effect T XO2 Ur
9.70 4.35 −1.85
850.53 171.05 30.94
S S S
18.28 −4.93 −5.28
1106.34 80.35 92.18
S S S
Interactions T ∗ XO2 T ∗ Ur XO2 ∗ Ur T ∗ XO2 ∗ Ur
−0.20 −1.00 0.05 −0.10
0.36 9.04 0.02 0.09
NS S NS NS
−1.48 0.18 −0.23 −1.08
7.21 0.10 0.17 3.83
S NS NS NS
S: Significant, NS: Not significant.
Fig. 8. Propane conversion, yield of aromatics, selectivity to COx and aromatics for the TZFBR as a function of oxygen mole percentage at reference values, see Table 1.
Fig. 9. Effect of height of the propane feed entry point upon propane conversion, selectivity to COx and aromatics in TZFBR at reference operating conditions.
percentage and a steady state was obtained for all the experiments. It can be observed from Fig. 9 that as the height ratio increases, the propane conversion and COx selectivity decrease. Since at lower height ratio more catalyst is in contact with propane, therefore more coke is produced which results in increasing propane conversion and COx selectivity. In addition, the results show that selectivity to aromatics achieves a maximum at hh/H = 0.5. These results could be explained with the aid of the previous kinetic study [34]. With decreasing height ratio below than 0.5 (then large reaction zone) more coke is formed on the catalyst by means of the aromatics which leads to a decreased selectivity to aromatics and to an increased COx selectivity, while by increasing height ratio above 0.5 less aromatics and coke were produced, because of lower catalyst contact with propane feed.
3.6. Effect of the relative velocity Propane conversion, selectivity to aromatics and COx as a function of relative velocity are shown in Fig. 10. Other variables are at their reference values. It can be seen that the propane conversion and selectivity to aromatics decrease with the increase in the relative velocity, because the gas contact time in the reaction zone and possibility of contact with the catalyst particles decreased. Also, increasing of the gas superficial velocity leads to increase in the bubble size which results in an increase in the mass transfer resistance between the gas in the bubbles and the catalytic emulsion phase. Increase in COx selectivity with rising of the relative velocity is shown in Fig. 10. The results reveal that an increase of the relative
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4. Conclusion For the first time, the propane aromatization over Zn/HZSM-5 catalyst in a two-zone fluidized bed reactor was accomplished. In the reactor used in this study, the reaction and regeneration zones were created by separating the oxygen and propane feeds. The results obtained in this work show that, by using a TZFBR, the catalyst deactivation as a result of coke deposition can be compensated by continuously regenerating the catalyst, thus obtaining a steady state operation in TZFBR. This work demonstrated that the yield of aromatics in the TZFBR is highest than in the conventional fluidized and fixed bed reactors at the steady state operation, where the coke formation leads to the catalyst deactivation in the conventional reactors in the absence of oxygen which results in decrease of the propane conversion and selectivity to aromatics with time. The effects of parameters such as: temperature, relative velocity, O2 mole percentage and their interactions were studied using two-level factorial design. Temperature was the main parameter affecting the two-zone fluidized bed reactor performance. Also, the effect of the height of propane feed entry point alongside the other aforementioned parameters was studied individually and the optimum experimental conditions were investigated. Thus, the TZFBR has several advantages over the conventional reactors regards to aromatics production: (a) Reaction and regeneration of catalyst occur in the same vessel. (b) The heat produced from coke burning in catalyst regeneration zone provides a part of the energy necessary for endothermic aromatization reaction. (c) By separating the oxygen and hydrocarbon fed the flammable mixture is avoided.
Fig. 10. Effect of relative velocity upon propane conversion, selectivity to COx and aromatics in TZFBR. Other variables are at their reference values.
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Fig. 11. Propane conversion, selectivity to COx and aromatics for the TZFBR as a function of temperature at reference operating conditions.
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3.7. Effect of the temperature The effect of the temperature upon the propane conversion and product selectivity is presented in Fig. 11. Since more coke was produced in the propane aromatization process at the higher temperatures, more oxygen was needed at these temperatures in order to achieve a steady-state condition in the TZFBR. Accordingly, these experiments were carried out with a constant input of propane and different amount of oxygen (1.5–3% O2 mole percentage). Based on the previous developed kinetic model for the propane aromatization process [34], because the dehydrogenation, cracking and aromatization reactions are endothermic, as the temperature was raised, the propane conversion and selectivity to aromatics increased, as could be expected. Increasing of the COx selectivity with temperature is due to burning more coke formed at the higher temperature. 203
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