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Development of light naphtha aromatization process using a conventional fixed bed unit
Catalysts in Petroleum Refining and Petrochemical Industries 1995 M. Absi-Halabi et al. (Editors) 9 1996 Elsevier Science B.V. All rights reserved.
455
D E V E L O P M E N T OF LIGHT NAPHTHA AROMATIZATION PROCESS USING A CONVENTIONAL F I X E D B E D UNIT S. F u k a s e a, N. lgarashi a, K. Kato b, T. N o m u r a c and Y. lshibashi c
a Petroleum Laboratory, Japan Energy Corporation, 17-35 Niizo-minami 3 chome, Toda, Saitama, 335, Japan b Engineering Department, Petroleum Refining Division, Japan Energy Corporation, 10-1 Toranomon 2 Chome, Minato-ku, Tokyo, 105, Japan c Mizushima Oil Refinery, Japan Energy Corporation, 2- 1 Ushiodori, Kurashiki, Okayama, 712, Japan ABSTRACT A new process of light naphtha aromatization, LNA process, has been developed. The process converts light paraffins containing high concentration of C4-C6 paraffins to aromatics. The development of a new catalyst having a long term stability enabled us to use a conventional fixed bed unit. Based on the results of fundamental and scale-up studies, Japan Energy Corporation has operated 2,250 BPSD demonstration plant and confirmed the good stability of the catalyst. 1. INTRODUCTION Aromatics are mainly produced through the catalytic reforming of These days, light hydrocarbons have become alternative feedstocks production. Several processes have been developed for this reaction: former [2] and Aroformer [3]. The conditions of these processes technology of catalyst regeneration, such as continuous regeneration regeneration, due to rapid catalyst deactivation.
heavy naphtha. for aromatics Cyclar [1], Zrequire special or swing type
The economics of light hydrocarbon aromatization processes does depend on the initial investment cost, mainly construction cost, and the price difference between the feedstock and aromatics. Today's construction cost of refinery processes is becoming expensive. Due to the massive construction cost and no expected widening in the feedstock/BTX price difference, the payout years of a construction cost would be lengthy. One solution of this problem is to develop a new aromatization process using a conventional fixed bed, thus avoiding the need to construct CCR type or swing type reactor unit. Currently in many refineries, conventional "semiregenerated type" heavy naphtha reformers have been replaced by CCR reformers. A number of these units are currently unused and available for
456 another use of light naphtha aromatization. The objective of the development of LNA process, thus, is to develop a new catalyst having extended stability which enables us to use conventional fixed bed reactors, minimizing initial construction cost. Under these circumstances, Japan Energy Corporation has been conducted extensive research on the development of a new aromatization catalyst that exhibits high activity and excellent inhibition of coke formation. Based on this fundamental research, an LNA demonstration plant with a capacity of 2,250 BPSD has been operated in 1994. This paper describes the features of the LNA process and its performance. 2. FUNDAMENTAL AND SCALE-UP STUDIES 2.1. Experimental 2.1.1 Microflow Reactor The reaction was carried out in a stainless steel microflow reactor. In each run, a 2 g portion of catalyst was placed in the reactor and heated to 520 ~ under a nitrogen stream. The nitrogen stream was replaced by a light naphtha vapor fed by a micro- plunger pump. The reaction was carried out at 520 ~ under a pressure of 3 k g / c m 2 G with a WHSV of 0.7 h -1. The products were analyzed periodically by gas chromatography. The properties of the feedstock are shown in Table 1. Table 1 Properties and components of the Feedstock Density (g / cm 3) 0.6591 Sulfur (ppm) <0.1 Nitrogen (ppm) <0.3 H20 (ppm) 13 Components (wt%) n-C5 32.3 i-C5 + C5 naphthene 17.8 n-C6 15.1 i-C6 + C6 naphthene 24.9 n-C7 2.1 i-C7 + C7 naphthene 5.4 n-C8 0 i-C8 0.1 Benzene + Toluene 2.3
2.1.2. Small-scale pilot plant Studies using a small-scale pilot plant was performed. The stainless steel tubular reactor containing 170 g of catalyst was heated by an electric heater. The system pressure was set by a back pressure regulator. The catalyst bed was heated to 500 ~ under a nitrogen stream. The nitrogen stream was then replaced by a
457 light naphtha vapor fed by a plunger pump. Reaction was carried out under a pressure of 3 kg/cm2G, reaction temperature of 500 ~ and WHSV of 0.7 h -1. The p r o d u c t s were collected, w e i g h e d and a n a l y z e d p e r i o d i c a l l y by gas chromatography.
2.1.3. Catalyst Metallosilicates, such as gallo-aluminosilicate (Ga-Al-silicate), and zincoaluminosilicate (Zn-Al-silicate) were synthesized according to procedure described in the paper [4]. Hydrothermal crystallization was carried out in a stainless steel autoclave at 160 ~ for 20 h. Resulting crystals were separated from the solution through a centrifuging process and washed with water. The samples were contacted with a 1M NH4NO3 solution at 80 ~ for 2 h, and thereafter taken from the solution and washed with deionized water. The products were dried at 120 ~ for 12 h and then calcined in air at 540 ~ for 3.5 h. Secondary ammonium-ion-exchange was performed for the calcined sample. The ammonium-ion-exchanged zeolites were dried at 120 ~ overnight, and then calcined at 540 ~ for 3.5 h to obtain H § exchanged zeolites. Zinco-aluminosilicate (molar ratio Si/Al=30, Si/Zn=100) synthesized according to above method was designated as Zn-Al-silicate (A). Zinco-aluminosilicate synthesized without template (molar ratio Si/Al=22, Si/Zn=169) was designated as Zn-Al-silicate (B). Zn-Al-silicate (A) was furthur dealuminated by our proprietary technique of pressurized steaming and was named as Zn-Al-silicate (C). H-ZSM-5 was prepared according to the same procedure described above. Zinc-exchanged ZSM-5 (Zn/H-ZSM-5) was prepared by contacting H-ZSM-5 described above with 0.035M Zn(NO3)2 solution at 80 ~ for 2 h.
2.2. Results and discussion 2.2.1. Stability of Various Catalysts Prepared by Different Methods Aromatization of light naphtha was carried out in the microflow reactor, under the condition stated above. The catalysts examined were Zn-Al-silicate (A), Ga-Al-silicate (molar ratio Si/A1 = 25, Si/Ga = 35), H-ZSM-5 (Si/A1 = 25) and Zn/ZSM-5 (Si/A1 =25, Si/Zn = 150). All the catalysts were calcined at 540~ for 3.5h before use. Change in conversion is shown in Fig. 1. The conversion of light naphtha is defined here by the following equation and is calculated on a carbon number basis: Conversion = (Products- AR0)/(1 - AR0) (1) where, AR0 = fraction of aromatics in the feed naphtha Products = (H2 + (C1 to C4) + (C5 and C6 olefine) + aromatics) in product. All of the catalysts demonstrated 100% conversion at very initial stage of the reaction u n d e r the condition employed. They showed, however, great differences in the stability of the activity. While H-ZSM-5 showed rapid decline of conversion, on the other hand, Zn-Al-silicate (A), showed more stable activity. Accordingly, metallosilicate exhibited better stability.
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2.2.2. Stability of Catalysts with Different Pore Volume and External Surface Area Zinco-aluminosilicates having different pore volume and external surface area were synthesized and their stability in aromatization was investigated. Table 2 shows some properties of catalysts. External surface area was measured by benzene-filled pore method described by Inomata et al. [5]. Average crystal size was measured geometrically by SEM images. Stability of those catalysts was examined under the same condition as stated above, and the results are shown in Fig. 2. Table 2 Properties of zinco-aluminosilicates Pore Volume (ml / g) Zn-A1-Silicate (A) Zn-A1-Silicate (B)
0.374 0.195
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Crystal Size (~m) 0.3 3
459 Zn-Al-silicate (B), having smaller pore volume and external surface area, demonstrated greater aging rate. Coke deposited on Zn-Al-silicate (B) was found to be less than that on Zn-Al-silicate (A). These suggest that greater deactivation of Zn-Al-silicate (B) was due to pore mouth blocking. Therefore, morphology of the zeolite is a key factor in maintaining stability of the catalyst.
2.2.3. Stability of Catalysts with Different Acidity Stability of zinco-aluminosilicates having different acidity was examined. Temperature Programed Desorption (TPD) of ammonia was carried out to measure the acidity according to the method reported by Niwa et al. [6]. Comparison of TPD spectra for two zinco-aluminosilicates are shown in Fig. 3.
Zn-AI-silicate (A)
A
100 200 300 400 500 Temperature (C) Fig. 3 NH3 - TPD spectra
Their stability in aromatization was also examined in the same way stated above. Zn-Al-silicate (C), having a smaller intensity of the TPD peak showed far greater stability than Zn-Al-silicate (A) as is shown in Fig. 4. The analysis of the amount of coke deposited on the catalyst showed 28 wt% for Zn-Al-silicate (A) after the experiment of 170 h and 44 wt% for Zn-Al-silicate (C) after 2000 h of experiment, indicating lower coke-forming rate of the latter catalyst. This indicates the importance of acid-property control in order to prepare the catalyst with a long-term stability. Typical yield pattern of Zn-Al-silicate (C) is also shown in Table 3.
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Table 3 Typical Yield Product H2 C1 - C2 C3 - C5 Benzene Toluene Xylene C9 + A r o m a t i c s
(wt%) 1.8 25.8 21.3 12.9 22.9 11.9 3.4
Stability of the catalyst after regeneration The stability of the catalyst after regeneration was examined using the smallscale pilot plant. A catalyst having a relatively higher coke-forming rate was used for this experiment to carry out the experiment in a short period. Regeneration was carried out at 450 ~ under a diluted air pressure of 5 k g / c m 2 G for 96 h. The variation of the conversion is shown in Fig. 5.
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461 No substantial decline of activity was found after regeneration. No change in chemical or physical properties of the catalyst was also observed before and after the regeneration. The catalyst showed high stability towards regeneration. 3. PROCESS DESIGN The concept of LNA process is to utilize an existing fixed bed unit, which is typically unused catalytic reformer replaced by CCR, with minor modification. Based on the results of the fundamental and scale-up studies, a 2,250 BPSD demonstration plant was designed, which was originally used as a conventional platforming unit.
3.1. Reactor Design A kinetic reaction model was developed and used with a process simulator to conduct reactor design [7]. The model was also used to investigate the optimum catalyst loading pattern to achieve the highest conversion and aromatics yield in the demonstration plant which consists of three reactors with preheaters in series. This was because a large temperature decrease was expected through the adiabatic reactors of the plant due to the endothermic nature of the whole reactions. 3.2. Preheater furnace tube design Adiabatic and endothermic reaction requires preheating and inter-heating furnaces in the reaction section. Even though hydrogen is being produced in the reaction, there is no hydrogen in the feed preheater furnace. In contrast, in most refinery furnace operations, when hydrocarbon vapor is heated higher than 500 ~ the vaporized feed usually contains hydrogen or steam to prevent coking inside heater tubes. Thermodynamic considerations [8] indicate that our reaction condition also is within the coking region which could cause tube coking. We have, therefore, chosen shorter residence time in the heater tubes to avoid severe coking during furnace operations. 3.3. Catalyst regeneration Small-scale pilot studies have demonstrated that the extended stability of the catalyst is such that it is well suited for service in semi-regenerative type operation using conventional fixed bed reactors. Semi-regeneration is a remarkable feature of this process compared with other light paraffin aromatization processes. Attention is required in regenerative operation to avoid catalyst degradation, which is due to coke-burning, because zeolite structure may collapse in severe hydrothermal atmosphere at higher temperature. In a commercial adiabatic reactor there is some concern about temperature distributions inside the reactor, in comparison with the isothermal small-scale pilot plant. We investigated, therefore, temperature profile in the catalyst bed to scale-up the process. The following study was conducted to investigate the temperature distributions during regeneration of large packed column.
462 Heat transfer of packed bed has been the subject of numerous studies. For cylindrical packed columns, a solution for determining temperature distributions was given using Bessel functions. Here, it is important to find out exact effective thermal conductivity of bed because of flowing gas and relatively high temperatures. Radial temperature distributions are more important than that of axial direction because the latter can be measured and controlled during the operation. Results of investigation have shown that the maximum temperature drop through radial direction does not exceed 45 ~ thus supporting system's operability [9].
4. PROCESS DESCRU~ION OF THE DEMONSTRATION PLANT A simplified flow diagram of the LNA demonstration plant is shown in Fig. 6. The plant is designed to prove the technology at semi-commercial scale. Process facilities consist of two major sections : reaction and product recovery. The reaction section includes a preheater, interheaters, and reactors. The recovery section includes separator, atmospheric fractionator. The predominance of dehydrogenation and cracking reactions causes the overall sequence to be highly endothermic. The LNA process employs adiabatic reaction stages with interheater to achieve optimum conversion and selectivity to aromatic product. The fresh feed, which consists of C4, C5, and C6, is directly charged into the preheater without combining with a recycle stream of unconverted feed. This is because high conversion level can be attained during the operation and almost no unconverted feed is obtained. Hydrogen stream from the separator, however, can be recycled to investigate the effect of hydrogen to hydrocarbon ratio at the inlet of the reactor.
R(
Gas
Heaters
LPG
Reactors
~eparazor
Reformate Stabilizer
~ ) Feedstock Fig.6 Simplified Flow Diagram of LNA Process
463 Semi-regeneration is a remarkable feature of this process compared with other light paraffin aromatization processes. A temperature swing adsorption type dryer is installed to eliminate water produced by coke-burning during regeneration, in order to avoid catalyst degradation. It is designed to remove water from the reactor effluent gas, thus avoiding water accumulation in the recycle gas stream. The method of regeneration is similar to that of conventional catalytic reforming. By introducing diluted air into the catalyst bed, the coke on the catalyst is burned off. This is carried out in several stages. The first stages is carried out under relatively mild conditions. In later stages, the temperature and the inlet oxygen concentration are increased. The temperature of the catalyst bed and the oxygen concentration at the outlet are monitored. Based on the basic design stated above, a demonstration plant was set up to verify the LNA process in Japan Energy Corporation's Mizushima Oil Refinery in Japan. The plant has been operated to aromatize light naphtha at the rate between 1,500 ~ 2,250 BPSD in 1994. 5. PROCESS PERFORMANCE OF DEMONSTRATION PLANT After the start-up, the plant achieved long-term operation without any trouble. Similar reaction yield acquired in the small-scale pilot plant was observed for several sets of operating conditions. The stability of the catalyst in the demonstration plant is shown in Fig. 7.
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464 6. CONCLUSION A new light naphtha aromatization process has been developed using a conventional fixed bed reactor. Fundamental study revealed the importance of preparation method, morphology, and acid property to increase the catalyst stability. Based on fundamental and scale-up studies, a demonstration plant was designed and operated. This operation confirmed the good stability of the catalyst. ACKNOWLEDGEMENT The demonstration plant work has been sponsored by Petroleum Energy Center in Japan, which is supported by Japanese Ministry of International Trade and Industry. The authors wish to thank Professor T. Inui of Kyoto University for his valuable suggestions. REFERENCES
1. C. D. Gosling, F. P. Wilcher, L. Sullivan and R. A. Mountford, Hydrocarbon Processing, 70 (12) (1991) 69. 2. S. Saito, K. Hirabayashi, S, Shibata, T. Kondo, k. Adachi and S. Inoue, Paper presented at 1992 NPRA annual meeting, New Orleans, March 22-24, 1992, AM-92-38 3. J. C. Barbier and A. Minkkinen, Paper presented at 1990 JPI Petroleum Refining Conference, Tokyo, October, 1990. 4. S. Fukase, H. Kumagai and T. Suzuka, Appl. Catal. A-General, 93 (1992) 35. 5. M. Inomata, M. Yamada, S. Okada, M. Niwa and Y. Murakami, J. Catal., 100 (1986) 264. 6. M. Niwa, M. Iwamoto and K. Segawa, Bull. Chem. Soc. Japan, 59 (1986) 3735. 7. K. Kato, S. Fukase, T. Amaya and Y. Sato, J. Japan Petrol. in press. 8. J. M. Harrison, J. F. Norton, R. T. Derricott and J. B. Mariot, Wekstoffe u. Korrosion, 30 (1979) 785. 9. K. Kato and S. Fukase, J. Japan Petrol. Inst. 37 (1994) 77.
455
D E V E L O P M E N T OF LIGHT NAPHTHA AROMATIZATION PROCESS USING A CONVENTIONAL F I X E D B E D UNIT S. F u k a s e a, N. lgarashi a, K. Kato b, T. N o m u r a c and Y. lshibashi c
a Petroleum Laboratory, Japan Energy Corporation, 17-35 Niizo-minami 3 chome, Toda, Saitama, 335, Japan b Engineering Department, Petroleum Refining Division, Japan Energy Corporation, 10-1 Toranomon 2 Chome, Minato-ku, Tokyo, 105, Japan c Mizushima Oil Refinery, Japan Energy Corporation, 2- 1 Ushiodori, Kurashiki, Okayama, 712, Japan ABSTRACT A new process of light naphtha aromatization, LNA process, has been developed. The process converts light paraffins containing high concentration of C4-C6 paraffins to aromatics. The development of a new catalyst having a long term stability enabled us to use a conventional fixed bed unit. Based on the results of fundamental and scale-up studies, Japan Energy Corporation has operated 2,250 BPSD demonstration plant and confirmed the good stability of the catalyst. 1. INTRODUCTION Aromatics are mainly produced through the catalytic reforming of These days, light hydrocarbons have become alternative feedstocks production. Several processes have been developed for this reaction: former [2] and Aroformer [3]. The conditions of these processes technology of catalyst regeneration, such as continuous regeneration regeneration, due to rapid catalyst deactivation.
heavy naphtha. for aromatics Cyclar [1], Zrequire special or swing type
The economics of light hydrocarbon aromatization processes does depend on the initial investment cost, mainly construction cost, and the price difference between the feedstock and aromatics. Today's construction cost of refinery processes is becoming expensive. Due to the massive construction cost and no expected widening in the feedstock/BTX price difference, the payout years of a construction cost would be lengthy. One solution of this problem is to develop a new aromatization process using a conventional fixed bed, thus avoiding the need to construct CCR type or swing type reactor unit. Currently in many refineries, conventional "semiregenerated type" heavy naphtha reformers have been replaced by CCR reformers. A number of these units are currently unused and available for
456 another use of light naphtha aromatization. The objective of the development of LNA process, thus, is to develop a new catalyst having extended stability which enables us to use conventional fixed bed reactors, minimizing initial construction cost. Under these circumstances, Japan Energy Corporation has been conducted extensive research on the development of a new aromatization catalyst that exhibits high activity and excellent inhibition of coke formation. Based on this fundamental research, an LNA demonstration plant with a capacity of 2,250 BPSD has been operated in 1994. This paper describes the features of the LNA process and its performance. 2. FUNDAMENTAL AND SCALE-UP STUDIES 2.1. Experimental 2.1.1 Microflow Reactor The reaction was carried out in a stainless steel microflow reactor. In each run, a 2 g portion of catalyst was placed in the reactor and heated to 520 ~ under a nitrogen stream. The nitrogen stream was replaced by a light naphtha vapor fed by a micro- plunger pump. The reaction was carried out at 520 ~ under a pressure of 3 k g / c m 2 G with a WHSV of 0.7 h -1. The products were analyzed periodically by gas chromatography. The properties of the feedstock are shown in Table 1. Table 1 Properties and components of the Feedstock Density (g / cm 3) 0.6591 Sulfur (ppm) <0.1 Nitrogen (ppm) <0.3 H20 (ppm) 13 Components (wt%) n-C5 32.3 i-C5 + C5 naphthene 17.8 n-C6 15.1 i-C6 + C6 naphthene 24.9 n-C7 2.1 i-C7 + C7 naphthene 5.4 n-C8 0 i-C8 0.1 Benzene + Toluene 2.3
2.1.2. Small-scale pilot plant Studies using a small-scale pilot plant was performed. The stainless steel tubular reactor containing 170 g of catalyst was heated by an electric heater. The system pressure was set by a back pressure regulator. The catalyst bed was heated to 500 ~ under a nitrogen stream. The nitrogen stream was then replaced by a
457 light naphtha vapor fed by a plunger pump. Reaction was carried out under a pressure of 3 kg/cm2G, reaction temperature of 500 ~ and WHSV of 0.7 h -1. The p r o d u c t s were collected, w e i g h e d and a n a l y z e d p e r i o d i c a l l y by gas chromatography.
2.1.3. Catalyst Metallosilicates, such as gallo-aluminosilicate (Ga-Al-silicate), and zincoaluminosilicate (Zn-Al-silicate) were synthesized according to procedure described in the paper [4]. Hydrothermal crystallization was carried out in a stainless steel autoclave at 160 ~ for 20 h. Resulting crystals were separated from the solution through a centrifuging process and washed with water. The samples were contacted with a 1M NH4NO3 solution at 80 ~ for 2 h, and thereafter taken from the solution and washed with deionized water. The products were dried at 120 ~ for 12 h and then calcined in air at 540 ~ for 3.5 h. Secondary ammonium-ion-exchange was performed for the calcined sample. The ammonium-ion-exchanged zeolites were dried at 120 ~ overnight, and then calcined at 540 ~ for 3.5 h to obtain H § exchanged zeolites. Zinco-aluminosilicate (molar ratio Si/Al=30, Si/Zn=100) synthesized according to above method was designated as Zn-Al-silicate (A). Zinco-aluminosilicate synthesized without template (molar ratio Si/Al=22, Si/Zn=169) was designated as Zn-Al-silicate (B). Zn-Al-silicate (A) was furthur dealuminated by our proprietary technique of pressurized steaming and was named as Zn-Al-silicate (C). H-ZSM-5 was prepared according to the same procedure described above. Zinc-exchanged ZSM-5 (Zn/H-ZSM-5) was prepared by contacting H-ZSM-5 described above with 0.035M Zn(NO3)2 solution at 80 ~ for 2 h.
2.2. Results and discussion 2.2.1. Stability of Various Catalysts Prepared by Different Methods Aromatization of light naphtha was carried out in the microflow reactor, under the condition stated above. The catalysts examined were Zn-Al-silicate (A), Ga-Al-silicate (molar ratio Si/A1 = 25, Si/Ga = 35), H-ZSM-5 (Si/A1 = 25) and Zn/ZSM-5 (Si/A1 =25, Si/Zn = 150). All the catalysts were calcined at 540~ for 3.5h before use. Change in conversion is shown in Fig. 1. The conversion of light naphtha is defined here by the following equation and is calculated on a carbon number basis: Conversion = (Products- AR0)/(1 - AR0) (1) where, AR0 = fraction of aromatics in the feed naphtha Products = (H2 + (C1 to C4) + (C5 and C6 olefine) + aromatics) in product. All of the catalysts demonstrated 100% conversion at very initial stage of the reaction u n d e r the condition employed. They showed, however, great differences in the stability of the activity. While H-ZSM-5 showed rapid decline of conversion, on the other hand, Zn-Al-silicate (A), showed more stable activity. Accordingly, metallosilicate exhibited better stability.
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2.2.2. Stability of Catalysts with Different Pore Volume and External Surface Area Zinco-aluminosilicates having different pore volume and external surface area were synthesized and their stability in aromatization was investigated. Table 2 shows some properties of catalysts. External surface area was measured by benzene-filled pore method described by Inomata et al. [5]. Average crystal size was measured geometrically by SEM images. Stability of those catalysts was examined under the same condition as stated above, and the results are shown in Fig. 2. Table 2 Properties of zinco-aluminosilicates Pore Volume (ml / g) Zn-A1-Silicate (A) Zn-A1-Silicate (B)
0.374 0.195
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Crystal Size (~m) 0.3 3
459 Zn-Al-silicate (B), having smaller pore volume and external surface area, demonstrated greater aging rate. Coke deposited on Zn-Al-silicate (B) was found to be less than that on Zn-Al-silicate (A). These suggest that greater deactivation of Zn-Al-silicate (B) was due to pore mouth blocking. Therefore, morphology of the zeolite is a key factor in maintaining stability of the catalyst.
2.2.3. Stability of Catalysts with Different Acidity Stability of zinco-aluminosilicates having different acidity was examined. Temperature Programed Desorption (TPD) of ammonia was carried out to measure the acidity according to the method reported by Niwa et al. [6]. Comparison of TPD spectra for two zinco-aluminosilicates are shown in Fig. 3.
Zn-AI-silicate (A)
A
100 200 300 400 500 Temperature (C) Fig. 3 NH3 - TPD spectra
Their stability in aromatization was also examined in the same way stated above. Zn-Al-silicate (C), having a smaller intensity of the TPD peak showed far greater stability than Zn-Al-silicate (A) as is shown in Fig. 4. The analysis of the amount of coke deposited on the catalyst showed 28 wt% for Zn-Al-silicate (A) after the experiment of 170 h and 44 wt% for Zn-Al-silicate (C) after 2000 h of experiment, indicating lower coke-forming rate of the latter catalyst. This indicates the importance of acid-property control in order to prepare the catalyst with a long-term stability. Typical yield pattern of Zn-Al-silicate (C) is also shown in Table 3.
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Table 3 Typical Yield Product H2 C1 - C2 C3 - C5 Benzene Toluene Xylene C9 + A r o m a t i c s
(wt%) 1.8 25.8 21.3 12.9 22.9 11.9 3.4
Stability of the catalyst after regeneration The stability of the catalyst after regeneration was examined using the smallscale pilot plant. A catalyst having a relatively higher coke-forming rate was used for this experiment to carry out the experiment in a short period. Regeneration was carried out at 450 ~ under a diluted air pressure of 5 k g / c m 2 G for 96 h. The variation of the conversion is shown in Fig. 5.
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400 600 800 1000 Time on Stream (h) Fig. 5 Catalyst Stability after Regeneration
461 No substantial decline of activity was found after regeneration. No change in chemical or physical properties of the catalyst was also observed before and after the regeneration. The catalyst showed high stability towards regeneration. 3. PROCESS DESIGN The concept of LNA process is to utilize an existing fixed bed unit, which is typically unused catalytic reformer replaced by CCR, with minor modification. Based on the results of the fundamental and scale-up studies, a 2,250 BPSD demonstration plant was designed, which was originally used as a conventional platforming unit.
3.1. Reactor Design A kinetic reaction model was developed and used with a process simulator to conduct reactor design [7]. The model was also used to investigate the optimum catalyst loading pattern to achieve the highest conversion and aromatics yield in the demonstration plant which consists of three reactors with preheaters in series. This was because a large temperature decrease was expected through the adiabatic reactors of the plant due to the endothermic nature of the whole reactions. 3.2. Preheater furnace tube design Adiabatic and endothermic reaction requires preheating and inter-heating furnaces in the reaction section. Even though hydrogen is being produced in the reaction, there is no hydrogen in the feed preheater furnace. In contrast, in most refinery furnace operations, when hydrocarbon vapor is heated higher than 500 ~ the vaporized feed usually contains hydrogen or steam to prevent coking inside heater tubes. Thermodynamic considerations [8] indicate that our reaction condition also is within the coking region which could cause tube coking. We have, therefore, chosen shorter residence time in the heater tubes to avoid severe coking during furnace operations. 3.3. Catalyst regeneration Small-scale pilot studies have demonstrated that the extended stability of the catalyst is such that it is well suited for service in semi-regenerative type operation using conventional fixed bed reactors. Semi-regeneration is a remarkable feature of this process compared with other light paraffin aromatization processes. Attention is required in regenerative operation to avoid catalyst degradation, which is due to coke-burning, because zeolite structure may collapse in severe hydrothermal atmosphere at higher temperature. In a commercial adiabatic reactor there is some concern about temperature distributions inside the reactor, in comparison with the isothermal small-scale pilot plant. We investigated, therefore, temperature profile in the catalyst bed to scale-up the process. The following study was conducted to investigate the temperature distributions during regeneration of large packed column.
462 Heat transfer of packed bed has been the subject of numerous studies. For cylindrical packed columns, a solution for determining temperature distributions was given using Bessel functions. Here, it is important to find out exact effective thermal conductivity of bed because of flowing gas and relatively high temperatures. Radial temperature distributions are more important than that of axial direction because the latter can be measured and controlled during the operation. Results of investigation have shown that the maximum temperature drop through radial direction does not exceed 45 ~ thus supporting system's operability [9].
4. PROCESS DESCRU~ION OF THE DEMONSTRATION PLANT A simplified flow diagram of the LNA demonstration plant is shown in Fig. 6. The plant is designed to prove the technology at semi-commercial scale. Process facilities consist of two major sections : reaction and product recovery. The reaction section includes a preheater, interheaters, and reactors. The recovery section includes separator, atmospheric fractionator. The predominance of dehydrogenation and cracking reactions causes the overall sequence to be highly endothermic. The LNA process employs adiabatic reaction stages with interheater to achieve optimum conversion and selectivity to aromatic product. The fresh feed, which consists of C4, C5, and C6, is directly charged into the preheater without combining with a recycle stream of unconverted feed. This is because high conversion level can be attained during the operation and almost no unconverted feed is obtained. Hydrogen stream from the separator, however, can be recycled to investigate the effect of hydrogen to hydrocarbon ratio at the inlet of the reactor.
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463 Semi-regeneration is a remarkable feature of this process compared with other light paraffin aromatization processes. A temperature swing adsorption type dryer is installed to eliminate water produced by coke-burning during regeneration, in order to avoid catalyst degradation. It is designed to remove water from the reactor effluent gas, thus avoiding water accumulation in the recycle gas stream. The method of regeneration is similar to that of conventional catalytic reforming. By introducing diluted air into the catalyst bed, the coke on the catalyst is burned off. This is carried out in several stages. The first stages is carried out under relatively mild conditions. In later stages, the temperature and the inlet oxygen concentration are increased. The temperature of the catalyst bed and the oxygen concentration at the outlet are monitored. Based on the basic design stated above, a demonstration plant was set up to verify the LNA process in Japan Energy Corporation's Mizushima Oil Refinery in Japan. The plant has been operated to aromatize light naphtha at the rate between 1,500 ~ 2,250 BPSD in 1994. 5. PROCESS PERFORMANCE OF DEMONSTRATION PLANT After the start-up, the plant achieved long-term operation without any trouble. Similar reaction yield acquired in the small-scale pilot plant was observed for several sets of operating conditions. The stability of the catalyst in the demonstration plant is shown in Fig. 7.
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400 600 800 1000 Time on Stream (h) Fig. 7 Stability of the Catalyst in the Demonstration Plant
464 6. CONCLUSION A new light naphtha aromatization process has been developed using a conventional fixed bed reactor. Fundamental study revealed the importance of preparation method, morphology, and acid property to increase the catalyst stability. Based on fundamental and scale-up studies, a demonstration plant was designed and operated. This operation confirmed the good stability of the catalyst. ACKNOWLEDGEMENT The demonstration plant work has been sponsored by Petroleum Energy Center in Japan, which is supported by Japanese Ministry of International Trade and Industry. The authors wish to thank Professor T. Inui of Kyoto University for his valuable suggestions. REFERENCES
1. C. D. Gosling, F. P. Wilcher, L. Sullivan and R. A. Mountford, Hydrocarbon Processing, 70 (12) (1991) 69. 2. S. Saito, K. Hirabayashi, S, Shibata, T. Kondo, k. Adachi and S. Inoue, Paper presented at 1992 NPRA annual meeting, New Orleans, March 22-24, 1992, AM-92-38 3. J. C. Barbier and A. Minkkinen, Paper presented at 1990 JPI Petroleum Refining Conference, Tokyo, October, 1990. 4. S. Fukase, H. Kumagai and T. Suzuka, Appl. Catal. A-General, 93 (1992) 35. 5. M. Inomata, M. Yamada, S. Okada, M. Niwa and Y. Murakami, J. Catal., 100 (1986) 264. 6. M. Niwa, M. Iwamoto and K. Segawa, Bull. Chem. Soc. Japan, 59 (1986) 3735. 7. K. Kato, S. Fukase, T. Amaya and Y. Sato, J. Japan Petrol. in press. 8. J. M. Harrison, J. F. Norton, R. T. Derricott and J. B. Mariot, Wekstoffe u. Korrosion, 30 (1979) 785. 9. K. Kato and S. Fukase, J. Japan Petrol. Inst. 37 (1994) 77.