Ind. Eng. Chem. Res. 2002, 41, 6015-6019
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Experimental Study of the Deep Catalytic Cracking Process in a Downer Reactor Rensheng Deng, Fei Wei,* Yan Jin, Qihao Zhang, and Yong Jin Department of Chemical Engineering, Tsinghua University, Beijing 100084, People’s Republic of China
A hot downer reactor of 4.5 m in height and 13 mm in inner diameter was built to examine its performance in the deep catalytic cracking (DCC) process. The experimental results show that the downer reactor can significantly improve the selectivity of desired products in comparison with the commercial riser adopting the same feed and catalyst: the yields of propylene and gasoline increase by 3.71 and 7.30 wt %, respectively, while the yields of dry gas and coke reduce by 6.77 and 1.98 wt %, respectively. The “close to plug flow” pattern of the gas and solids in the downer makes it a desirable reactor for the DCC process. Introduction The most widely used process for light olefins production is thermal cracking in the tubular furnace, which provides over 95% ethylene in the whole world. However, the coke produced in the cracking process deposits on the furnace inner wall and blocks up the reactor; thus, the suitable feeds for this process are limited to light hydrocarbons such as ethane, naphtha, and kerosene. Novel processes dealing with heavy feeds are appealed to meet the ever-increasing demands for light olefins. Deep catalytic cracking (DCC) is such a process developed by SINOPEC in recent years, treating heavy feeds such as vacuum gas oil and deasphalted oil.1 This process is colicensed by Stone & Webster, and seven commercial plants are running in China and Thailand. It has two operation modes: DCC-I works at the temperature of 560 °C, with the desired product of propylene; DCC-II works at the temperature of 530 °C, with the desired product of isoolefins.2 However, because of the severe operating conditions adopted, the riser reactor used in this process reveals a severe problem: the extensive backmixing of gas and solids results in high yields of low-value byproducts, the dry gas and coke, which can be up to 11.74 and 9.39 wt %, respectively, in the total products.3 The downer reactor was proposed by some famous companies such as Mobile,4,5 Texaco,6 and StoneWebster7 in order to overcome the shortcomings of the riser. A lot of research has been done in the study of hydrodynamics and mixing behavior in downer reactors, as well as the inlet and outlet.8-12 Hot models were also tested, and exciting results were reported. UOP developed a MSCC reactor for the fluidized catalytic cracking (FCC) process from which the gasoline yield is 6.6 vol % higher than that from the riser.13 Ikeda and Ino14 studied a downer FCC pilot plant for light olefin production and obtained a higher gasoline yield and a lower dry gas yield compared to the riser. However, little work has been done in the DCC process, where the influence of the reactor on the product distribution is more important than that for the FCC process. * Corresponding author. E-mail:
[email protected]. cn.
Figure 1. Experimental apparatus.
In this paper, a hot model of a downer reactor is built to examine its performance in the DCC process. Experimental Apparatus The experimental apparatus is shown in Figure 1. The air from the compressor flows through a pressure swing adsorption system, where the nitrogen is separated and sent to the catalyst heater and the stripper serving as the fluidization gas. The fluidized catalyst in the heater is heated to 700 °C and then enters the mixer through a catalyst feeder. At the same time, the feed from the tank is pumped, preheated to 350 °C, and sprayed into the mixer through two opposite injectors. With the help of the atomizing steam, the liquid feed is split into droplets of about 20-60 µm outside the ejectors, which then mix and contact with the catalyst. The cracking reaction occurs on the hot catalyst particles during the downflowing along the reactor. At the end of the downer, the gas and solids are separated in a quick separator. The catalyst, having lost its catalytic activity because
10.1021/ie010731n CCC: $22.00 © 2002 American Chemical Society Published on Web 10/31/2002
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Ind. Eng. Chem. Res., Vol. 41, No. 24, 2002 Table 1. Mass Balance
feed products gas
Figure 2. Real boiling point of the feed.
of the deposition of the coke on the activity site, falls down into a stripper where the products absorbing on the particles are stripped by steam. The products gas, together with the steam and fluidization gas, passes through a quench and two sequential coolers where the liquid products are condensed and separated. The liquid is collected and weighed; the gas is measured with a rotameter and sampled for analysis. The whole apparatus is about 12 m high, and the top temperature limit of the catalyst heater is about 1200 °C. The reactor is 4.5 m in height and 13 mm in inner diameter, with a feedstock rate of 3.5 kg/h. Feed, Catalyst, and Products The feed used in the experiments is deasphalted oil. It is a black solid at room temperature and turns into a sticky liquid when heated to above 70 °C. Its density is 0.8972 kg/m3, and the average molecular weight is 360. The concentrations of hydrogen, carbon, sulfur, and nitrogen are 12.85wt%, 86.08wt%, 0.56wt%, and 1525.24 ng/µL, respectively. Figure 2 shows the real boiling point of the feed. The main fractions are between 350 and 550 °C.
Figure 3. Chromatogram of the gas mixture.
weight (kg/h)
wt % of the feed
2.76 2.66 1.54
100.00 96.33 55.60
liquid coke loss
weight (kg/h)
wt % of the feed
0.94 0.18 0.10
34.20 6.53 3.67
The catalyst used is the CHP-1 zeolite catalyst. To obtain the highest yield of propylene, the catalyst is designed to have high catalytic activity for cracking reactions and low activity for hydrogen-transfer reactions. Its packing density is 0.84 kg/cm3, and the average diameter of the particles is 59 µm. The gas products are mainly composed of hydrogen and hydrocarbons of low carbon number. The hydrogen is analyzed by a gas chromatograph (SQ-206) with a thermal conductivity detector, and the hydrocarbons are analyzed by a gas chromatograph (Shimadzu 14B) with a flame ionization detector. The chromatograph chart of the hydrocarbons mixture is shown in Figure 3. The corresponding peaks are identified as methane, ethane, ethylene, propane, propylene, etc. The liquid products contain a lot of heavy hydrocarbons that cannot be easily analyzed one by one. The ASTM distillation test method is applied to classify the liquid into gasoline (BP ) 30-200 °C), light gas oil (LGO; BP ) 200-350 °C), and cycle stock (BP ) >350 °C). The coke depositing on the catalyst particles is analyzed with a coke analyzer. The sampled deactivated catalyst is heated to 700 °C in the flow of air; thus, the coke is burned and completely converted into carbon dioxide with the help of a hot cupric oxide powder. The carbon oxide is then analyzed by a gas chromatograph, and the content of coke on the catalyst can be measured. Table 1 lists the typical quantity relationship of feed and products in one run. A summary of 20 groups of data shows that the loss of mass balance is less than 5 wt %. Operating Conditions Scheme Because the hydrodynamics and mixing behavior of the downer are different from the riser, the operating
Ind. Eng. Chem. Res., Vol. 41, No. 24, 2002 6017 Table 2. Comparison of the DCC Process between the Commercial Riser and the Hot Downer Operating Conditions
feeding temperature reaction temperature operating pressure catalyst/oil ratio residence time
units
commercial riser
°C °C MPa kg/kg s
380 565 0.08 10.87-11.80 3.0-4.0
downer experiment 350 606 0.008 37.54 0.74
Material Balance name feed products Figure 4. Effect of the cracking temperature on the products yield.
Figure 5. Effect of the catalyst-to-oil ratio on the products yield.
conditions should also be adjusted to meet the demands of cracking reactions occurring in the downer reactor instead of copying those adopted in the riser. Because both the gas and solids flow in the direction of gravity, the catalyst will be accelerated to a higher velocity and thus a shorter contact time can be obtained compared to the riser operating at the same gas velocity and solids reflux. Usually, it will be less than 1 s in the downer. The higher particle velocity also results in a lower solids concentration in the downer, which reduces the conversion under the same operating conditions adopted in the riser. Thus, a higher reaction temperature, as well as a higher catalyst-to-oil ratio, is necessary to increase the reaction rate to achieve the desired conversion. As shown in Figure 4, the yields of methane, ethylene, and propylene in the downer DCC process increase with an increase of the cracking temperature. However, the methane and ethylene, which are the components of the dry gas, rise more significantly than propylene. In fact, the high yield of the dry gas in the commercial DCC riser has been a drawback for the wide application of the DCC process. As a result, the cracking temperature should not be too high. Figure 5 shows that an increase of the catalyst-to-oil ratio will lead to an increase of methane, ethylene, propylene, and butylene. However, different from temperatures, the propylene and butylene increase more significantly than the methane and ethylene. The different effects of temperature and the catalystto-oil ratio on the yields of the products can be explained by the mechanism of cracking reactions. Two reaction
1. dry gas H2S H2 CH4 C2H6 C2H4 2. LPG C3H8 C3H6 C4H10 C4H8 3. gasoline 4. LGO 5. coke 6. loss summary
commercial riser 100 11.03 0.20 0.29 3.39 2.14 5.01 39.77 3.27 19.54 2.84 14.12 21.47 19.23 8.5 0 100
downer experiment 100 4.26 / 0.13 0.91 0.36 2.86 42.97 0.89 23.25 3.46 15.37 28.77 17.82 6.52 3.67 100
paths coexist in the DCC process: one is the carbonium ion mechanism taking place at the activity site of the zeolite catalyst particles (catalytic cracking), by which the bulk propylene and butylene are produced, without hydrocarbons of less than two carbon atoms;15 the other is the free-radical mechanism occurring mainly in the gas phase (thermal cracking), which is influenced significantly by the temperature.16 Raising the cracking temperature will increase the thermal cracking remarkably and result in a sharp increase of methane and ethylene, while a large catalyst-to-oil ratio will enhance the catalytic effect (that is, the carbonium mechanism) and result in a significant increase of propylene and butylene. In summary, compared to the riser, a proper higher temperature, shorter contact time, and larger catalystto-oil ratio are desirable for the DCC process in the downer reactor. Results and Discussion Typical Experimental Results. The 15 000 tons/ year commercial DCC plant in the Jinan Refinery Factory (a subsidiary of SINOPEC) is the first commercial plant for the DCC process that was put into operation in 1994. The feed for the plant is the deasphalted oil mentioned above, and the catalyst is the CHP-1 catalyst. The cracking reactor is a riser with a fluidized bed on its top. The operating conditions and products distribution are shown in Table 2. For the purpose of comparison, typical data from the hot downer reactor are also listed in Table 2. As shown in Table 2, the operating temperature in the downer reactor is 41 °C higher than that of the riser, and the ratio of catalyst to oil is 2-3 times of that the commercial reactor. The residence time is 0.7-0.8 s, only one quarter of that in the riser. The gas is not necessary to suspend the particles in the downer, which
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Ind. Eng. Chem. Res., Vol. 41, No. 24, 2002
Figure 6. Comparison of products distribution between the riser and downer.
results in a lower pressure than that in the riser and benefits the feed conversion because cracking is a molecule-increasing reaction. The dry gas from the downer DCC process is suppressed significantly in comparison with that of the commercial riser: its yield is 4.26 wt %, about 38% of that of the riser. The yield of liquefied petroleum gas (LPG) is 3.20 wt % higher than that of the riser, among which the propane reduces while the propylene increases by 3.71 wt %. In the liquid products, the yield of gasoline is 28.77 wt %, 7.30 wt % higher than that of the riser, while the LGO is 17.82 wt %, 1.41 wt % lower. The total coke yielded on the catalyst drops from 8.50 wt % in the riser to 6.52 wt % in the downer. The products in ascending sequence of molecular weight are listed as shown in Figure 6: the lightest product, dry gas, decreases by 60% from the riser; the second, LPG, increases by 12%; the third, gasoline, increases by 39%; the LGO decreases by 4%; the heaviest, coke, decreases by 20%. It can be concluded that the products with very high and very low molecular weight (dry gas, LGO, and coke) are reduced significantly in comparison with the riser reactor, while the products with middle molecular weight (liquid gas and gasoline) increase remarkably. The products tend to concentrate on those with middle molecular weight in comparison with that of the riser, which indicates that the selectivity of desired products is improved in the downer reactor. This is desirable for the cracking process obviously. Influence of the Operating Conditions. Because the operating conditions in the downer are different from those in the riser, a question will arise: Is the improvement of selectivity caused by the optimal operating conditions or caused by the close to plug flow pattern in the downer reactor? As shown in Figures 4 and 5, increases of the cracking temperature and the catalyst-to-oil ratio will lead to an increase of propylene. Thus, higher temperature and larger C/O ratio adopted in the downer surely contribute to the high yield of propylene. However, according to Figures 4 and 5, methane and ethylene will also increase at severe cracking conditions, which conflicts with the fact that the dry gas is relatively low in the experiments. Therefore, the improvement of the selectivity appearing in the hot downer DCC process cannot be explained only by the variation of operating conditions from the riser. Desired Properties of the Downer Reactor. In the riser reactor, both the gas and solids flow against the direction of the gravity, which results in the severe nonuniformity in the radial flow structure and large
axial backmixing.17,18 The clustering of particles near the wall forms a dense catalyst phase, where the reaction rate is high; on the contrary, the dilute core region with high gas velocity causes the relatively low reaction rate. As a result, the feed near the wall tends to overcrack and turns into dry gas and coke, while the feed in the core cannot convert fully. Additionally, the large axial backmixing of gas and solids, which is followed with a wide residence time distribution, will also lead to significant secondary reactions and bring about high yields of byproducts but low yields of desired products. Different from the riser, the cocurrent downflow of the gas and solids along the gravity direction eliminates the mechanism for particles clustering so that the radial distribution of the particle velocity, gas velocity, and solids concentration is uniform. The axial dispersion coefficient is about 1-2 orders of magnitude less than that of the riser, which leads to a very narrow residence time distribution in the downer.11,12 When the mean residence time of the feed meets the optimal reaction time for the propylene yield, the selectivity of propylene will be increased and a high yield will be obtained. The simulation carried out by Yasemin et al.,19 Wei et al.,20 and Deng et al.21 on the downer FCC process shows that both the less axial backmixing and uniform radial flow structure are responsible for the high selectivity of the desired products in the downer. It is the nature of close to plug flow pattern in the downer that results in the desired products distribution in the cracking processes. Scale-Up Effect. Another question is about the scaleup effect in the downer application. Usually, the effect is negative for commercial reactors because of the nonideal flow pattern and other problems. However, the hydrodynamics are quite similar in downers of different diameters.9,22-24 Furthermore, two aspects in the commercial reactors are more desirable than those in the pilot plant: one is the wall effect, which is much less in the reactors of large diameter; the other is the contact efficiency between the gas and solids at the mixer. In the pilot plant, the volume of the mixer is so limited that the mixture of feed droplets and catalyst particles cannot expand fully; thus, some of the droplets may fall down on the wall, and some may collide with each other and aggregate into big drops. The mixing of droplets and solids may be insufficient in the narrow space. All of these can damage the yields of olefins and will be avoided or abated in large-scale reactors. Hence, the conclusion that the commercial plants will give out lower yields of desired products cannot be drawn here. Additionally, the improvement of the feed-catalyst contact efficiency in commercial plants will be of help to reduce the apparent catalyst-to-oil ratio. For example, an efficiency of 50% means that half of the particles cannot contact the feed and accelerate the cracking reaction rate; that is, they are “wasted” although being counted in the apparent catalyst-to-oil ratio. The redesign and optimization of the mixing part is an important task for the commercial application of the downer, and special attention should be focused on this topic. Conclusions As the first attempt to apply the downer reactor to the DCC process, the experimental results from this work show a promising products distribution: dry gas and the coke are suppressed significantly in comparison
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with the commercial riser with the same feed and catalyst, while the desired propylene and gasoline are raised. The analysis of the process shows that it is the excellent properties of the downer reactor that lead to the significant improvement of selectivity. Commercial experiments in larger-scale reactors are to be carried out for the application of the downer DCC process. Acknowledgment The financial assistance from the Natural Science Foundation of China under Contract Nos. 29736190 and 29725613 is acknowledged. Literature Cited (1) Li, Z.-T.; Jiang, F.-K.; Min, E.-Z., et al. Production of Light Olefins by DCC Process. Shiyou Lianzhi Yu Huagong 1989, 7, 3133. (2) Picciotti, M. Novel Ethylene Technologies Developing, But Steam Cracking Remains King. Oil Gas J. 1997, 95 (25), 53-58. (3) Yu, B.-D.; Shi, Z.-C.; Xu, Y.-H., et al. The Commercial Application of CRP-1 Zeolite and Operation of the 150,000 t/a DCC Plant. Shiyou Lianzhi Yu Huagong 1995, 26 (5), 7-12. (4) Gross, B. Heat Balance in FCC Process and Apparatus with Downflow Reactor Riser. U.S. Patent 4,411,773, Oct 5, 1983. (5) Gross, B.; Ramage, M. P. Fluid Catalytic Cracking, esp. for Heavy OilssUsing Downflow Riser Reactor and Fast Regenerator. U.S. Patent 4,385,985, May 31, 1983. (6) Bunn, D. P.; Niccum P. K. Improved Fluid Catalyst Hydrocarbon Conversion ProcesssWhere Hot Regenerated Catalyst Meets Hydrocarbon Feedstock in Gravity Flow Catalytic Reactor Followed by Ballistic Separator. U.S. Patent 4,514,285, Apr 30, 1995. (7) Gartside, R. J. Olefin Production from Heavy Hydrocarbon Feed. U.S. Patent 4,663,019, May 5, 1987. (8) Bai, D.-R.; Jin, Y.; Yu, Z.-Q.; Gan, N.-J. Gas-solids Flow Patterns in a Concurrent Downflow Fast Fluidized Bed (CDCFB). J. Chem. Ind. Eng. (China) 1991, 6 (2), 171-181. (9) Wang, Z.; Bai, D.; Jin, Y. Hydrodynamics of Concurrent Downflow Circulating Fluidized Bed (CDCFB). Powder Technol. 1992, 70, 271-275. (10) Cao, C.-S.; Jin, Y.; Yu, Z.-Q.; Wang, Z.-W. The Gas-solids Velocity Profiles and Slip Phenomenon in a Concurrent Downflow Circulating Fluidized Bed. In Circulating Fluidized Bed Technology IV; Avidan, A. A., Ed.; AIChE: New York, 1994; pp 406413.
(11) Wei, F.; Wang, Z. W.; Jin, Y.; et al. Dispersion of Lateral and Axial Solids in a Concurrent Downflow Circulating Fluidized Bed. Powder Technol. 1994, 81, 25-31. (12) Wei, F.; Jin, Y.; Yu, Z.-Q.; et al. Gas Mixing in the Cocurrent Downflow Circulating Fluidized Bed. Chem. Eng. Technol. 1995, 18, 59-62. (13) Kauff, D.; Bartholic, D.; Steves, C.; et al. Successful Application of the MSCC Process. NPRA Annual Meeting, 1996. (14) Ikeda, S.; Ino, T. Down-Flow FCC Pilot Plant for Light Olefin Production. Proceedings of Fifth International Symposium on the Advances in Fluid Catalytic Cracking Presented before the Division of Petroleum Chemistry, Inc., 218th National Meeting, New Orleans, LA, Aug 1999; American Chemical Society: Washington, DC, 1999; pp 487-489. (15) Corma, A.; Planelles, J.; Sanchez, J.; et al. The Role of Different Types If Acid Site in the Cracking of Alkanes on Zeolite Catalysts. J. Catal. 1985, 93 (1), 30-37. (16) Albright, L. F.; Crynes, B. L.; Corcoran, W. H. Pyrolysis: Theory and Industrial Practice; Academic Press: New York, 1983. (17) Berruti, F.; Chaouki, J.; Godfroy, L.; et al. Hydrodynamics of Circulating Fluidized Bed Risers: A Review. Can. J. Chem. Eng. 1995, 73 (Oct), 579-602. (18) Wei, F.; Cheng, Y.; Jin, Y.; et al. Axial and Lateral Solids Dispersion in a Binary-Solid Riser. Can. J. Chem. Eng. 1998, 76 (1), 19-26. (19) Yasemin, G.; Pugsley, T. S.; Berruti, F. Computer Simulation of the Performance of Fluid Catalytic Cracking Risers and Downers. Ind. Eng. Chem. Res. 1994, 33, 3043-3052. (20) Wei, F.; Ran, X.; Zhou, R.; et al. Dispersion Model for Fluid Catalytic Cracking Riser and Downer Reactors. Ind. Eng. Chem. Res. 1997, 36, 5049-5053. (21) Deng, R.-S.; Wei, F.; Liu, T. F.; Jin, Y. Radial Behaviors in Riser and Downer during FCC Process. Chem. Eng. Process. 2002, 41 (3), 259-266. (22) Zhang, H.; Zhu, J.-X.; Bergougnou, M. A. Hydrodynamics in Downflow Fluidized Beds (1): Solids Concentration Profiles and Pressure Gradient Distributions. Chem. Eng. Sci. 1999, 54, 54615470. (23) Lehner, P.; Wirth, K.-E. Characterization of the Flow Pattern in a Downer Reactor. Chem. Eng. Sci. 1999, 54, 54715483. (24) Zheng, Y.; Qian, Z.; Wei, F.; et al. Modeling of Hydrodynamics in Downer Reactor. Proceedings of Third Joint China/USA Chemical Engineering Conference (CUChE-3); Tsinghua University Press: Beijing, 2000; Vol. 12, pp 97-106.
Received for review September 4, 2001 Revised manuscript received May 31, 2002 Accepted May 31, 2002 IE010731N
