Article pubs.acs.org/EF
Laboratory-Scale Experiments and Industrial Practice of LowTemperature Contact and High Ratio of Catalyst to Oil in the FCC Process Gang Wang,* Zekun Li, Yongfeng Li, Jinsen Gao,* Chunming Xu, Yongmei Liang, and Xiaoqin Wang State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Beijing, 102249, China ABSTRACT: Contrast experiments of fluid catalytic cracking (FCC) were conducted in a technical pilot scale riser FCC apparatus. The results showed that the use of a low-temperature regenerated catalyst and high ratio of catalyst to oil (CTO) improved the initial contact between the catalyst and feedstock. Thermal cracking was inhibited and catalytic cracking was promoted in the FCC riser, which were favorable for increasing the conversion and generation of desirable products. A modified FCC process was proposed and numerical simulation were used in the selection of the equipment structure for mixing cold and hot regenerated-catalyst streams to achieve optimal contact between the catalyst and the feedstock. A commercial unit was revamped based on the experimental results to test the modified FCC process. Results from industrial practice showed that product distributions under optimal conditions of low-temperature contact and high ratio of catalyst to oil in the modified FCC process were superior to those from routine FCC at nearly similar reaction conditions. The modified FCC process yielded an increase in total feed conversion of about 9 wt %; small decreases in the coke and dry gas yields were also obtained.
1. INTRODUCTION Fluid catalytic cracking (FCC) is the most important conversion process used by petroleum refineries to convert high-boiling hydrocarbon fractions of petroleum crude oils into more valuable products, such as gasoline, and olefinic gases, among others. As of 2006, over 400 petroleum refineries operate FCC units worldwide and about one third of the crude oil refined in those refineries is processed through FCC to produce high-octane gasoline and fuel oils.1 Typical feeds that are processed are gas oils derived from atmospheric and vacuum fractionators. As a consequence of the crude oil supply becoming even heavier in its contents, FCC feeds are increasingly more refractory to crack. However, economic pressures encourage the refineries to process higher boiling range oils, such as atmospheric residue (AR), vacuum residue, and severely solvent-extracted oils to gain higher profit. Especially in China, the blending ratio of residue oils (boiling range of over 550 °C) is beyond 35 wt %.2 Residue FCC (RFCC) is a main process in China. Therefore, refiners continuously have to change and adjust their process conditions to improve product distributions. Novel ways of operating commercial FCC units include short contact time catalytic cracking,3 millisecond catalytic cracking,4,5 and two-stage residue catalytic cracking.6 Shorter contact time units produce greater catalyst to oil (CTO) ratios and allow the reaction to occur at higher temperatures. In these processes, a high regenerated-catalyst temperature is considered to improve the heat transfer from the regenerated catalyst to the liquid feed and enhance thermal shocks for the shattering of Conradson carbon residue (CCR) or asphaltenes and other heavier unstable molecules. The effect of thermal shock is focused on the conversion of CCR. However, an overly high temperature enhances the formation of hydrogen and C1−C2 molecules. Recently, RxCat (formerly X-design)7−10 and a process for residue catalytic cracking and gasoline upgrading11 © 2013 American Chemical Society
were proposed, the important characteristic of which are a lower regenerated-catalyst temperature and greater CTO ratio. A lower regenerated-catalyst temperature and a higher circulation rate of catalyst enhance catalytic cracking making it more efficient than thermal cracking. Thus, dry gas yield is reduced and light oil yield is increased. The influence of CTO ratio, temperature, and residence time scale on coke formation, as well as the reaction behavior and the kinetic model processing of vacuum gas oil (VGO) feedstock, have been intensively investigated in laboratory-scale once-through microriser reactors in which the residence time can be varied by changing the reactor length without changing the catalyst and feedstock flow rates. Such systemic works have provided significant insights into the catalytic cracking behavior and mechanism of VGO.12 However, many problems are encountered during catalytic cracking, and such decreases in the temperature of the catalyst with increasing CTO ratio. To the author’s best knowledge, no report has yet specifically addressed these problems. Whereas dry gas is the least valuable product of catalyst cracking, containing mostly materials with rich hydrogen, it still exerts serious effects on the hydrogen balance and yields of the objective products. For decarbonization processes such as RFCC, dry gas and coke are the ultimate products of heavy oil to lighter fuel conversion, and these products have very high contents of hydrogen (Table 1 shows the typical hydrogen contents of FCC products) and carbon, respectively. The RFCC reaction is a hydrogen autobalance reaction system, and the key to obtaining good product distribution is to transfer as much hydrogen and carbon as possible from the feedstocks to the intermediates such as gasoline and diesel. Therefore, reducing the yield of dry Received: October 17, 2012 Revised: February 3, 2013 Published: February 5, 2013 1555
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Table 1. Typical Hydrogen Content in FCC Products item
data (wt.%)
dry gas (H2, C1 and C2) LPG (C3 and C4) gasoline diesel slurry coke
>19 >15 13.2−13.8 10.2−12.2 9.8−10.5 6.1−6.7
Table 3. Properties of the Feedstock item
gas and coke is important for the redistribution of hydrogen and carbon during production. In this article, the performance of a modified RFCC system utilizing low regenerated-catalyst temperature and large CTO ratio conditions was investigated in a technical pilot scale riser (TPSR) FCC apparatus and compared with the results of a system utilizing conventional conditions. By analyzing the mechanism of dry gas and coke formation, a preferential FCC process for lower dry gas and coke yield was proposed. A commercial CFD software called FLUENT 6.3.26 was applied to simulate the structure design of the prelifting stage and obtain optimal contact between the catalyst and feedstock. To assess the proposed FCC process, a commercial unit was revamped according to the experimental and numerical simulation results, and the operating conditions and distribution of products were investigated.
Table 2. Physicochemical Properties of the Equilibrium FCC Catalyst Item
Data 67.0 0.31 95.2 0.85
item
0.9130 47.43
group analysis/wt.% saturate
3.30 577 86.77
hydrogen/wt % sulfur/wt % nitrogen/wt % nickel/(ug/g) vanadium(ug/g)
12.87 0.12 0.23 4.6 0.2
aromatics resin + asphaltene boiling point composition/°C IBP/5% 10/30% 50/70% 81.5%
data 57.08 27.61 15.31
302/384 420/475 515/543 568
in the reaction. Using the TPSR, the cracking of a paraffinicbased AR feed was studied at residence times varying from 0.6 to 3.5 s with little change in the flow regime in the riser. The main difference between the TPSR and the commercial FCC unit is that the TPSR can be operated at adiabatic or isothermal conditions by electrical surface heat, which eliminates heat transfer through the vessel walls. The catalyst circulation rate was adjusted by controlling the pressure difference between the regenerator and the stripper and is thus not limited by heat balance as in the operation of a commercial FCC unit. The TPSR was controlled using a computer, and the main operation parameters can be shown and adjusted at any time during the experiments. During all of the tests, the output measured was between 97 wt % and 100 wt % of the injected feed, which is indicative of reasonably good mass balances. In this work, the conversion is defined as the amount of feed converted into dry gas, liquefied petroleum gas (LPG), gasoline, and coke. 2.3. Operating Parameters. 2.3.1. CTO Ratio. To simulate the commercial industry unit, the CTO ratio was calculated through the heat balance at the riser inlet. A schematic of the local heat balance at the riser inlet is shown in Figure 2. A good thermal insulation layer was provided at the riser inlet, thus heat loss was omitted. We simplify the heat balance with a basic assumption that negligible heat from reaction has been incurred prior to measurement of mix temperature. The heat quantity carried into the riser, Qin, can be calculated by
2. EXPERIMENTAL SECTION 2.1. Catalyst and Feedstock. A commercial, rare-earth modified ultra stable zeolite-based equilibrium FCC catalyst was used in the laboratory-scale experiments. The catalyst had a fairly high pore volume and specific surface and a large concentrated particle size distribution. The physicochemical properties of the catalyst are listed in Table 2. The feed used in the experiments was a paraffinic-base feedstock representative of those used in RFCC units in China; its detailed properties are shown in Table 3.
Microactivity index Pore volume (cm3/g) Surface area (m2/g) Packing density (g/cm3) Particle size distribution (wt.%) 0 to 40 μm 40 to 80 μm >80 μm
data
density (20 °C)/(g/cm3) refractive index (80 °C)/ (mm2·s−1) CCR/wt % molecular weight carbon/wt %
Q in = Ccat(T1 − T2)Gcat
(1)
where Ccat is the specific heat of the catalyst and equals 1.100 kJ/kg·°C. T1 and T2 represent the regenerated-catalyst temperature and the mix temperature (°C), respectively, and Gcat is the catalyst circulating rate (kg/h). The heat quantity absorbed by the incoming fluids, Qout, can be calculated by
13 55 32
Q out = Q oil + Q H O + Q N
2.2. Experimental Setup. The laboratory-scale experiments were performed in a TPSR FCC apparatus, which can be operated with continuous cracking reaction and catalyst regeneration, similar to a commercial FCC unit. The outline of the equipment is shown in Figure 1. The TPSR is a pilot scale entrained flow reactor in which the reaction time can be varied by changing the location of the feeder, thereby also changing the reactor length. Traditionally, with a pilot scale reactor, the residence time can be changed by adjusting the flow rate in the riser, which would have a significant influence on the flow regime. However, the reaction time cannot vary over a wide range. Obviously, the flow regime plays a large role
2
2
(2)
where Qoil, QH2O, and QN2 are the heats absorbed by the oil, atomizing steam, and prelifting nitrogen respectively the values of which can be calculated from their respective inlet temperatures and the mix temperatures. For example, Qoil can be calculated by Q oil = (H2 − Hin)(T2 − Tin)Goil
(3)
where Tin is the inlet temperature of oil, H2 and Hin represent enthalpies of the oil at T2 and Tin respectively, and Goil is the oil feeding rate (kg/h). 1556
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Figure 1. Schematic representation of the technical scale pilot plant.
2.3.3. Reaction Temperature. In the riser reactor, the reaction temperature is defined as the temperature of gas and oil at the exit of the riser reactor. In the current study, the temperature of the regeneration catalysts is defined as the temperature of the catalyst in the inclined regeneration tube. 2.4. Product Analysis. The cracked gas was measured by a flowmeter and analyzed using an Agilent 6890 gas chromatograph with three detectors. ChemStation software was used to determine the fractions of H2, N2, and C1 to C6 hydrocarbons in the gas products. The coke generated from the reactions in the pilot plant was determined by a CO and CO2 analyzer at the flue gas outlet. The flue gas volume was measured by a flowmeter. The liquid product was weighed after collection and analyzed by simulated distillation using another Agilent 6890 gas chromatograph according to the ASTM-2887-D method. The cuts were made at 204 °C for gasoline, 350 °C for diesel, and higher than 350 °C for heavy cycle oil (HCO). The group composition of the feed and the products were analyzed separately as feed, gas, gasoline, diesel, and HCO fraction. The gas fraction was divided into H2, paraffins, and olefins. The gasoline fraction was determined by PIONA analysis and divided into paraffins, olefins, naphthenes, and monocyclic aromatics.
Figure 2. Schematic of the local heat balance at the riser inlet.
Qin equals Qout. Thus, by obtaining the value of Gcat from the heat balance of the unit, CTO can be calculated by CTO =
Gcat Goil
(4)
2.3.2. Residence Time. For the TPSR unit, the feedstock was injected at the bottom of the riser and then came into contact with the high-temperature catalysts from the inclined regeneration tube. The residence time of the gas and oil in the riser can be described as the ratio of the length of the riser reactor to the mean velocity of the gas and oil.
3. RESULTS AND DISCUSSION 3.1. Laboratory-Scale Experiments of FCC. The performance of an FCC unit, as well as its product distribution and selectivities, such as byproducts (coke and dry gas) and desirable products (gasoline and diesel) in the reaction system, is dependent on a large number of parameters. Aside from the feed composition and catalyst properties, the residence time, CTO ratio, and temperature influence the conversion process in their own ways. The effects of these parameters on the performance of residue catalytic cracking were investigated through four experimental conditions. The first experiment was performed under standard operation conditions to simulate a typical commercial industry unit, which was labeled Base. The experimental conditions were as follows: regenerated-catalyst temperature, 690 °C; CTO ratio, 6; mix temperature of oil and
Residence time = length of riser reactor /mean velocity of gas and oil
The mean velocity of the gas and oil in the riser is: U=
Uout − Uin ln(Uout /Uin)
(4)
where Uin and Uout refer to the velocity of the steam, nitrogen, and gasified feedstocks at the entrance and exit of the riser reactor (m/s), respectively. Their values can be obtained from their input quantity and the sectional area of the riser reactor. 1557
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catalyst, 530 °C; and residence time ranging from 0.6 to 3.0 s. Other conditions were set to investigate the effect of regenerated-catalyst temperature and CTO ratio. These conditions are listed in Table 4.
catalyst surface. The heavy components contain the main contributors for feedstock carbon residue. These contributors are unable to diffuse into the catalyst pore during cracking reactions due to their larger molecule size and more aromatic nucleus. Highly condensed contributors were converted into additive coke by thermal condensation reactions in the initial oil catalyst mixing step, except for some cracks in the long side chains. The carbon residue of the feedstocks used was 3.30%, and the coke yields in the first 0.6 s of the initial mixing stage were between 4.50% and 7.5% for all of the tests in Figure 3. The coke yield reached as high as 7.5% when the CTO ratio was 15 due to the increase in initial conversion. On the basis of a previous study, 40% to 80% of the carbon residue of the feedstocks can be converted into coke.10 Thus, the coke formed in the initial mixing stage from an increase in conversion was obtained from not only additive coke but also catalytic coke. Conversion of as much carbon residue as possible into products can reduce the content of additive coke; thus, the regeneratedcatalyst temperature and CTO ratio should match each other. From Figure 3, we can observe that the content of coke on catalysts can be reduced under larger CTO ratio resulting in the increase of catalytic activity during reaction, and thus the conversion increased in the same reaction time. However, much higher CTO ratio also can strengthen the decarbonization of feedstock, and therefore increase the yield of coke. Especially, when the combination of higher CTO ratios with the regenerated-catalyst temperatures are not suitable, there are two consequences. One is that the high boiling point components cannot be vaporized sufficiently lacking of the heat at a low regenerated-catalyst temperature, and then increase the coke yield due to increased formation of liquid coke; another is that abundant thermal cracking can take place and then lead to excessive dry gas and coke at a high regenerated-catalyst temperature. Therefore, the regeneratedcatalyst temperature should be adjusted according to the properties of feedstock to realize the sufficient vaporization of feedstock and high catalytic selectivity in the reactor. Catalytic coke is a final product of aromatics obtained through bimolecular hydrogen transfer reactions;14 thus, in the later part of the riser reactor, the secondary reactions of olefins and aromatics induce more catalytic coke. As shown in Figure 3, the yields of coke increase in the following cracking stage. The increased coke in this stage mainly consists of catalytic coke and the increase in coke yields is in accordance with that of conversion. Part a of Figure 3 also shows that a larger CTO ratio results in a lower coke content in the catalyst, indicating
Table 4. Operating Conditions of the Laboratory-Scale Experiments item
base
I
II
III
CTO regenerated-catalyst temperature (°C) mix temperature (°C) reaction temperature (°C) reaction time (s)
6 690
8 670
10 640
15 630
530
536
533
545
500
500
500
520
0.6−3.0
0.6−3.0
1.1−3.0
0.6−3.0
On the basis of the experiment results, a modified residue catalytic cracking process was proposed. The modified process was used in commercial practice according to the optimal reaction conditions. 3.1.1. Coke Yield versus Reaction Time. In the FCC process, coke, an essential product in the decarbonization process, can be divided into several types based on its formation mechanisms:13 catalytic coke, additive coke, contaminant coke, and catalyst to oil coke. Catalytic coke is dependent on the feedstock and catalyst, as well as on the residence time, which is associated with acid-catalyzed cracking reactions. Additive coke depends on the composition of the feedstocks, such as the Conradson carbon percentage, asphaltene content, and presence of multiring high molecular weight aromatic molecules. Contaminant coke is produced from heavy metal poisoning during the catalytic cracking reaction. Catalyst to oil coke is a measurement of unstripped but potentially strippable hydrocarbons. In our study, catalysts with a low level of metal content underwent sufficient stripping. Thus, the coke content can be considered as the sum of the catalytic coke and additive coke contents. Changes in the coke yields with reaction time under the three FCC operational conditions are presented in Figure 3. The coke yields increased sharply in the initial mixing stage and then increased slowly in the following cracking stage. In the initial mixing stage, the feedstock was injected into the riser and came into contact with the catalysts. The lighter components vaporized immediately whereas the heavy components failed to vaporize upon contact with the hot
Figure 3. Coke content of the catalyst (a) and coke yield (b) as a function of reaction time. 1558
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Figure 4. Dry gas yields and selectivity as a function of reaction time.
(part a of Figure 4) with increasing reaction times. The selectivity of dry gas slightly decreased with the increase in CTO ratio (part b of Figure 4). However, the effects of CTO ratio on dry gas yield should be taken into consideration when increasing it. Corma19 supposed that alkanes can be transformed into dry gas such as hydrogen, methane, ethane, and ethylene through monomolecular cracking reactions through forming penta-coordinated carbonium ions by hydrogen proton attacks. At early reaction times in the riser, hydrocarbons easily produce carbonium ions promoting catalytic cracking reactions because the catalysts are not covered with much coke and have relatively high activities. Therefore, different optimal CTO ratios for the reduction of dry gas apply to different catalysts and feedstocks. On the basis of the analysis, the reduction in dry gas requires optimization of the mixing conditions of gas and oil to reduce the thermal cracking reaction of the hydrocarbons on hot catalysts and improve catalytic activities in the following cracking stage, which can retard the thermal cracking reaction of the hydrocarbons in the later part of the riser due to the deactivation of catalysts. 3.1.3. Overall Conversion and Other Product Yields. Figure 5 shows that compared with the traditional case (regeneratedcatalyst temperature of 690 °C and CTO ratio of 6), case I (regenerated-catalyst temperature of 670 °C and CTO ratio of 8), case II (regenerated-catalyst temperature of 640 °C and CTO ratio of 10), and case III (regenerated-catalyst temperature of 630 °C and CTO ratio of 15) showed different
higher catalytic activity. Thus, a corresponding higher yield of coke can be obtained with the same reaction time (part b of Figure 3). From the analysis above, the reactions of nonvaporizing heavy components in the initial mixing stage should be optimized to obtain lower coke yields. Reaction conditions to vaporize and crack the heavy components should also be elucidated. Reducing the hydrogen transfer reactions of the aromatics in the oil to gas cracking stage is important to limit coke yields. The optimal operational conditions include a CTO ratio of 8 and a regeneration temperature of 670 °C for cracking poor-quality heavy oil in the riser based on our experiment. These conditions favor the cracking of residue carbon, reducing the production of additive coke, and avoiding overly high catalyst coke contents due to a higher CTO ratio in the later reaction. 3.1.2. Dry Gas Yield versus Reaction Time. In the initial mixing stage, a thermal-homolytic scission reaction in the C−C covalent bond occurs upon contact of the hot regenerated catalyst (with a temperature as high as 690 °C) with the hydrocarbons.15 Then, the free radicals formed continue to undergo beta scission and form ethylene and other free radicals.16 Scission continues until a methyl-free, ethylene-free, or hydrogen-free radical is formed, which can combine with the hydrogen-free radical from other hydrocarbons to produce methyl, ethylene, or hydrogen respectively and form dry gas. Dry gas is formed mainly in the initial mixing stage due to the thermal shock effects of the hot catalyst on heavy oil.17,18 In Figure 4, when the reaction time was 0.6 s, the dry gas yield was between 1.5 wt % and 2 wt %. When the reaction time was nearly 3.0 s, the dry gas yield increased in range from 3.5 wt % to 4.5 wt %. On the basis of the results, 44% of dry gas in total dry gas was produced in the initial mixing stage. In the following cracking stage, dry gas yield increased with increasing reaction time. The temperature was as high as 500 °C in the riser during heavy oil catalytic cracking, which is equal or higher than the temperature of heavy oil thermal cracking or delayed cracking. Catalytic cracking and thermal cracking are two competing reaction mechanisms in the riser due to the lower catalytic activity produced by the coke. As the reaction proceeds, the thermal reaction increases in the later part of the riser speeding up the formation rate of dry gas.12 To inhibit the formation of dry gas, the average catalytic activity in the reaction zone in the riser can be effectively increased by increasing the CTO ratio in the riser. Figure 4 shows that the yields of dry gas under all of the tested conditions increased
Figure 5. Conversions of feed as a function of reaction time. 1559
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distributions during conversion. With decreasing regeneratedcatalyst temperatures and increasing CTO ratios, conversion increased under the same reaction time. When the regeneratedcatalyst temperature decreased to 670 °C and the CTO ratio increased by two units, the conversion improved from about 55 wt % to 65 wt % within a reaction time of 0.6 s. The decrease in regenerated-catalyst temperature and increase in CTO ratio resulted in a shortened reaction time to obtain the same conversion. The diesel, gasoline, and LPG yields as a function of reaction time in different operation cases are shown in Figures 6,7, and
Figure 8. LPG yield as a function of reaction time.
reaction time. Comparing the four curves, we can observe that a higher yield of LPG can be reached under a lower regeneratedcatalyst temperature with higher CTO ratio. The yield of LPG was only about 8 wt % for the base experiment, whereas the LPG yield increased to about 15 wt % for case III within the first 0.6 s when the regenerated-catalyst decreased. Analyzing the distribution of dry gas and coke in the riser at different reaction times, three methods had been taken into consideration to reduce the yields of dry gas and coke under the same conversion: proper reduction of the temperature between the feedstock and regenerated catalysts to inhibit the thermal cracking reaction in the initial mixing stage; application of a high CTO ratio to improve the total reactive activity of the catalysts and control the thermal cracking reaction in the following cracking stage; and strengthening of the collision contact among catalyst particles, as well as between catalyst particles and heavy oil atomization droplets, which can effectively convey the heat brought by the regenerated catalysts to the heavy oil molecules and promote its vaporization and cracking. From the data above, we found that case I has a coke advantage relative to case II and case III, and also has a minimum content of dry gas throughout. Therefore, a modified FCC process was proposed in the following paragraph. 3.2. Industrial Practice of a Modified FCC Process. 3.2.1. Description of a Modified FCC Process. The CTO ratio, regeneration temperature, reaction temperature, and reaction time are interrelated in the industrial FCC unit and the coordination of these variables is determined by the thermal balance. Thus, adjustment of thermal balance between the riser and regenerator will enhance the flexibility of the FCC unit and result in higher economic effects. Some technologies for cooling down the regenerated catalyst have been developed and granted patents. However, no further progress has been made. On the basis of the discussion above, the combination of a suitable regenerated-catalyst temperature and a higher CTO ratio resulted in improved production distribution. Thus, a modified FCC process was put forward in our study to meet the operation conditions. The charts of the ordinary and modified FCC processes are shown in Figure 9. Compared with the ordinary FCC process, a modified FCC process has more flexibility in terms of reaction depth control. A method for the rapid separation of gas and the catalyst and a cooler for the cold catalyst were developed. Rapid separation of gas and the catalyst is necessary to control the reaction depth when the CTO ratio is increased. The revamped FCC process has an additional tube for cold catalysts and a
Figure 6. Diesel yield as a function of reaction time.
Figure 7. Gasoline yield as a function of reaction time.
8, respectively. Figure 6 shows that at the initial time, the diesel yield increased as the regenerated-catalyst temperature decreased and CTO ratio increased. In case III, at a regenerated-catalyst temperature of 630 °C and a CTO ratio of 15, the diesel yield nearly reached its maximum at 0.6 s. These results are in agreement with serial kinetics. A lower regenerated-catalyst temperature and higher CTO ratio seemed to result in a shorter reaction time to reach maximum diesel yields. Diesel is prone to cracking into gasoline and gas fractions if higher catalytic activities are provided. The yields of gasoline shown in Figure 7 indicate a trend similar to that of diesel. The yields of gasoline in the four cases first increased and then decreased indicating that secondary reactions occurred with increasing reaction time. Unlike the trends of diesel and gasoline, Figure 8 shows that the yields of LPG obtained from the four cases increased with the increase in 1560
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Figure 10. Different preriser structures for catalyst mixing.
Using the CFD method, detailed flow fields and temperature distributions in the prelifting stage of various structures were obtained, providing necessary information for the optimization of these prelifting structures. The temperature distributions relevant to the various prelifting structures are shown in Figure 11. Significantly nonuniform distribution of the temperature was observed in the prelifting structures shown in parts d and e of Figure 12. The temperature was high on the left side and low on the right side, which is in line with the solid distribution of the hot and cold catalysts in the reactor. In these structures, the hot catalyst was introduced from the left side, whereas the cold catalyst entered from the right. Consequently, the temperature showed nonuniformity in the radial direction of the reactor. In the structures shown in parts b and c of Figure 12, nonuniform distributions of the temperature were attenuated, although a lower temperature appeared in the region near the cold catalyst inlet. On the basis of temperature profiles, the first two structures discussed are superior to the latter two. Average temperature differences between the hot and cold catalysts at the outlet are illustrated in Figure 12. Figure 12 indicates that the structure in part b of Figure 10 exhibited the highest temperature difference (54 °C), whereas the structure in part c of Figure 10 showed the lowest temperature difference (20 °C). Therefore, the latter structure is superior to the former. The above numerical results indicate that the structure shown in part c of Figure 10 benefits from the mixing and heat transfer between the hot and cold catalysts leading to a lowtemperature difference at the outlet. Therefore, the structure in part c of Figure 10 was selected and used to revamp the industrial FCC unit. 3.2.3. Operating Conditions and Product Distribution. A comparison of operating parameters and product distributions before and after the industrial FCC unit revamp is shown in Table 5. We can observe that the properties of the feedstock after the revamp worsened, showing a larger density, a larger CCR, a higher sulfur content, and a poorer composition of saturate, aromatic, resin, and asphaltene. The product distribution after the industrial FCC unit revamp showed much better results. The total feed conversion increased from 63.84 wt % to 72.52 wt %. The gasoline yield increased by over 12 wt % and the coke and dry gas yields slightly decreased. 3.2.4. Properties of LPG, Gasoline, and Diesel. The properties of the FCC products before and after the industrial FCC unit revamp were analyzed, and the results are listed in Tables 6, 7, and 8. Table 6 shows the properties of LPG. The
Figure 9. Diagram of the FCC unit before and after revamp.
formal inclined tube for the hot regenerated catalysts connected to the riser, which allows the separate control of cold and hot catalyst recirculation mass and the flexible operation of regenerated-catalyst temperature and CTO. Unfortunately, this setup poses problems in terms of the uniform mixing of the two flows of cold and hot catalysts, as well as the avoidance of insufficient evaporation of heavy oil. Therefore, some parameters of the prelifting stage, such as the mixing point and the contact time after mixing, should be taken into consideration. 3.2.2. Numerical Simulations for the Optimal Catalyst Mixing Structure. The structures of the prelifting stage and the inlet location of the cold catalyst play an important role in the mixing and heat transfer between the hot and cold catalysts, which dominates the final temperature distribution in the reactor. Today, detailed experimental investigations on the prelifting stage of industrial riser reactors remain extremely challenging due to complex hydrodynamics and the lack of appropriate measuring and testing techniques. Computational fluid dynamics (CFD) has become a useful approach in exploring complex processes in the petrochemical industry with increasing computational power. In this work, a commercial CFD software, FLUENT 6.3.26, was applied to optimize the structure design of the prelifting stage. Various potential structures of the prelifting stage are shown in Figure 10. Part a of Figure 10 shows the conventional structure of the prelifting stage. The hot catalyst is introduced into the reactor from the left side and then entrained by the rising vapor. Part b of Figure 10 shows the new structure design of the reactor, with a sleeve located at the entrance of the cold catalyst. The hot and cold catalysts are introduced from the left and right sides of the reactor, respectively. In the structure shown in part c of Figure 10, both the hot and cold catalysts are introduced from the same side of the reactor. The other parts are similar to those in part b of Figure 10. The structure in part d of Figure 10 is identical to that in part c of Figure 10, except for the sleeve. In part e of 10, a catalyst mixer is located at the base of the prelifting stage. The hot and cold catalysts are introduced into the mixer from two different sides. 1561
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Figure 11. Temperature distributions in different preriser structures.
Table 5. Comparison of Operating Parameters and Product Distributions Before and After the Industrial FCC Unit Revamp item properties of feedstock density (at 20 °C) (g/cm3) refractive index (at 80 °C)/(mm2·s−1) CCR (wt %) sulfur content (μg/g) group analysis (wt.%) saturate aromatics resin asphaltene operating parameters temperature at riser outlet (°C) C/O reaction time/s temperature of regenerated catalyst (°C) ratio of HCO to fresh feedstock product distribution (wt.%) acid gas dry gas LPG gasoline diesel slurry coke loss feed conversion (wt.%)
Figure 12. Temperature differences at the outlet of different preriser structures.
total olefin content increased by over 2 wt % after the revamp. In particular, the butane content increased from 23.45 wt % before revamp to 26.34 wt % after revamp. The iso-butane content also showed an increasing trend after the revamp, which suggests the rapid separation of gas and catalyst. The low regenerated-catalyst contact temperature weakened the coking tendency of the catalysts, and enhanced secondary reactions, such as the isomerization reaction. The higher CTO ratio also improved the distribution of LPG increasing the content of total olefins. Table 7 shows a comparison of gasoline properties obtained before and after the industrial FCC unit revamp. After the revamp, the gasoline properties changed significantly. The density and sulfur contents increased and the RON slightly changed. Differences in the composition of gasoline relied on the contents of paraffins and olefins. From the diesel properties before and after the industrial FCC unit revamp listed in Table 8, remarkable changes can be observed. The cetane number and basic nitrogen content decreased. The density flash point, condensation point, nitrogen content, and sulfur content also increased after the revamp due to the higher catalyst activity and resultant overcracking.
before revamp
after revamp
905.5 18.76 3.62 3342
920.6 23.3 3.93 5111
65.38 21.35 12.28 0.99
54.8 23.3 21.4 0.61
501 6.0 3.32 685 0.2
505 7.65 3.48 671 0.08
0.45 3.11 17.11 34.53 30.39 5.77 8.64 0.35 63.84
0.47 2.91 14.15 46.54 24.29 2.89 8.45 0.30 72.52
Table 6. Comparison of LPG Compositions Before and After the Industrial FCC Unit Revamp
1562
item
before revamp
after revamp
total C2 propane propylene iso-butane N-butane butene total olefins
0.16 12.48 43.26 15.13 5.52 23.45 66.71
0.00 9.07 42.53 17.89 4.17 26.34 68.87
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Table 7. Comparison of Gasoline Properties Before and After the Industrial FCC Unit Revamp item
before revamp
after revamp
708.0 388 91.5
733.9 586 91.6
42.4 36.5 21.1
44.7 33.7 21.6
density (g/cm3) sulfur content (μg/g) RON composition of gasoline (v%) paraffins olefins aromatics
Table 8. Comparison of Diesel Properties before and after the Industrial FCC Unit Revamp item
before revamp
after revamp
density (g/cm3) flash point (°C) condensation point (°C) cetane number nitrogen content (μg/g) basic nitrogen content (μg/g) sulfur content (μg/g)
887.8 56.0 −7.0 34.0 1015 165 2656
938.6 81.5 −0.17 27.7 1460 130 5298
4. CONCLUSIONS Contrast experiments of FCC reactions showed that a lower regenerated-catalyst temperature and higher catalyst to oil ratio improve the initial contact condition between the catalyst and feedstock resulting in an increase of conversion and more desirable products. On the basis of the results, a modified FCC process was proposed. The numerical simulation results showed that the new structure design of the reactor, with a sleeve located at the entrance of the cold catalyst and where the hot and cold catalysts are introduced from the same side of the reactor, is the optimal structure to improve the product properties and reduce the yields of dry gas and coke. Industrial practice experiments were conducted in an FCC unit where the rapid separation of gas and catalyst and a cooler for the cold catalyst were developed. The optimal structure was utilized with a lower regenerated-catalyst temperature and a high CTO ratio resulting in increased conversion and decreased coke and dry gas yields.
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AUTHOR INFORMATION
Corresponding Authors
*Tel.: 8610-8973-3085, fax: 8610-6972-4721, e-mail:
[email protected] (G.W.). *Tel.: 8610-8973-3993; fax: 8610-6972-4721; e-mail: jsgao@ cup.edu.cn (J.G.). Notes
The authors declare no competing financial interest.
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ACKNOWLEDGMENTS The authors acknowledge the financial support provided by National Natural Science Foundation of China (21176252) and Key Technologies Research and Development Program of China (2012BAE05B02).
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REFERENCES
(1) Jones, D.S.J.; P. P., Pujado Handbook of Petroleum Processing: Springer; 2006. (2) He., M. Y. The development of catalytic cracking catalysts: Acidic property related catalytic performance. Catal. Today 2002, 73, 49−55. 1563
dx.doi.org/10.1021/ef301690u | Energy Fuels 2013, 27, 1555−1563
