Article pubs.acs.org/EF
Cross-Flow Naphtha Reforming in Stacked-Bed Radial Reactors with Continuous Solid Circulation: Catalyst Deactivation and Solid Circulation between Reactors Roberto Galiasso Tailleur* Department of Thermodynamics, Simon Bolivar University, Sartenejas, Baruta, Miranda, Venezuela Department of Chemical, Biological, and materials Engineering , Oklahoma University , Norman, Oklahoma 73019, USA S Supporting Information *
ABSTRACT: A naphtha-reforming reactor system, consisting of four stacked-bed reactors with cross-flow of gas and circulation of catalyst, was simulated. Commercial and micropilot plant data were used to adjust the kinetics and deactivation constants, based on previous results obtained in fixed-bed operations with another catalyst. The simulation of radial commercial reactors was performed using mixing zones for the intervessel solid transfer and a second zone through which the solids move downward and within which reactions take place. The mass and energy differential equations were solved in the radial and axial direction point by point, using the Runge−Kutta−Fehlberg (RKF) numerical method. The effects of solid mixing on catalyst profiles of coke were determined in a cold model. Sequential deactivation and partial mixing tests were performed to verify their effects on catalyst activities. Spent catalysts were characterized and tested under different operating conditions. The feed and reformate were characterized using an extended PIONA-MS method to obtain additional information about the isomers for different carbon numbers. Temperature, yields, activities and pressure are predicted by the model and compared to commercial data. The results show the importance of considering the effect of solid mixing, the two-dimensional pressure-drop temperature, and the variation in catalyst activity across and along the reactors. decreases from the first reactor to the fourth reactor at constant solid throughput. One generalized section of the reforming stacked-bed reactor (CFCCR) is shown schematically in Figure 1. This figure depicts (half) four stacked beds connected by a cone−transfer pipe−inverted cone device. The gas flow is indicated there by yellow arrows and the temperature profile is indicated by red dashed lines. The catalyst particles move downward under gravity in a vertical channel formed by two coaxial perforated cylinders, while the gas phase of the reactant mixture flows radially inward across the catalyst bed from an external perforated cylinder (distributor; for example, see detail in ref 17)) to another internal perforated cylinder (collector). The reactor system is composed of several stacked gas-phase reactors of either equal or increasing volume, in series. There are two key problems in modeling this type of reactor system. First is the manner in which the model takes into account the lack of uniform distribution of the gaseous reactant mixture along the axial height in the multiple catalyst beds, and another one is the blending of catalyst particles with different coke content and still-active sites between stages. Gas flow directly affects the movement of the particle bed due to the convergent momentum of the gas phase, with respect to solids, which would tend to produce “cavitation”18−20 in a gas flowrate above a certain value. The drag force on the particles from the gas flow results in a frictional
1. INTRODUCTION Radial gas flow continuous moving-bed cross-flow catalytic radial reactors (CFCCRs) boast numerous advantages, including low pressure drop, high flux capacity, the option of using small-sized catalysts, and easy coupling of chemical reaction and catalyst regeneration.1 The typical licensors of the continuous moving catalyst naphtha-reforming processes to produce high research octane number (RON) gasoline and aromatics are UOPs and IFPs. With continuous regeneration, the coke formed in the reactor is continuously burned off; thus, the catalyst achieves as-new performance when returned to the upper reactor.2 Although reforming plants are important components of modern refineries, and recent publications have discussed the simulation of this process,2−12 most published works regarding modeling and optimization of the naphtha-reforming process address the use of simplified kinetics and fluid dynamics. In addition other applications for radial reactors were recently published.13−15. The moving-bed radial reactor is particularly suitable for catalytic reactions in which the catalyst deactivated as the particle moves down, but operating continuously under pseudo-steady-state conditions, as a function of the time-onstream. The deactivated catalyst is withdrawn from the bottom of the reactors and the regenerated catalyst introduced back into the top of the reactors to continuously operate the system. The circulation rate of catalyst between stacked vessels is controlled by gravity using a special internal device (see ref 16, among others) to connect the vessel and avoid the axial flow of gas. The residence time of the solids in each of the vessels © 2012 American Chemical Society
Received: June 12, 2012 Revised: August 28, 2012 Published: August 28, 2012 6938
dx.doi.org/10.1021/ef300996z | Energy Fuels 2012, 26, 6938−6959
Energy & Fuels
Article
results obtained with a cold model in our group of research,31 partially included here, indicated that solid fluid dynamics can be considered using inlet and outlet solid backmixing zone (represented as zone A in Figure 1) connected by an intermediate plug-flow zone (zone B in Figure 1) to simulate the activity, selectivity, and stability of catalyst in commercial reactors. To support commercial operation of the naphtha reforming technology, solid transfers between vessels have been studied by the author by sampling and characterizing the catalyst during the operation and after the shutdown and by cold modeling. Deactivation is another important parameter to consider in the reforming design. Many papers have been published on the modeling of catalyst deactivation due to coking. Voorhies32 pioneered the correlation of coke deposition with time-onstream. In the 1980s, the author proposed a simulation using an empirical equation as a function of coke content.33 Since then, several authors have related catalyst activity to the amount of coke deposited on the catalyst (see refs 4−7, among others). Wolf and Peterson34 developed a kinetic expression for coke formation, and Beltramini et al.35 considered the deactivation of both the acid and metal functions of the catalyst. Several authors,29,30 and the author of this paper,8,36 started to use the concept of activity developed by Levenspiel.37 We found, during the microplant studies of commercial reforming catalyst, that the deactivation of acid sites on a catalyst, which is a slow event using modern catalysts, was effectively dependent on alkyl-cyclopentane and on the inverse of hydrogen partial pressure, temperature, level of activity, and time-on-stream; the deactivation of metal sites was only dependent on temperature, time-on-stream, and level of activity.8,33,36,38 Metal sites deactivation is only dependent on temperature. In CFCCRs, the catalyst suffers a progressive deactivation and coke accumulation while it moves downward in the reactor, producing changes in the reaction rate profile and, consequently, changes in radial gas flow rate and temperature. Previous fluid-dynamics studies in cold models22−25,31 demonstrated that, in the concentric cylindrical zone (Figure 1, zone B), the catalyst moves as a quasi-piston flow with small variation in bed density if the adequate linear velocity of gas is used. No information is published about the effect of at-rest time, bed settling, and partial solids mixing in the transfer zones Ai (see Figure 1) between vessels. In zone B, the catalyst located near the distributor works at higher temperatures than those near the collector, because of a decreasing profile of temperature product of endothermic reactions. These temperature profiles generate an increasing profile of activities in the direction of gas flow (centripetal operation with higher coke deposition and catalyst deactivation at the inlet and lower one at the outlet). In addition, as the catalyst moves downward, it is exposed to more and more reactant, resulting in additional coke accumulation and catalyst deactivation along the axis. At the outlet of zone B (bottom of the reactor), the catalyst has profiles of coke content, site activity, and temperature before being discharged to the next vessel through a cone− transfer line−cone-type device (zone Ai). When the solid fall into the next reactor, the particles are partially mixed, changing the profile of coke, activity, and temperature in the top of the subsequent vessel. In a commercial plant, the rate of solids circulation is controlled by gravity forces, friction forces, and a valve located at the outlet of the last reactor that seals the vessels and produces some discontinuity in the delivery of solids. Understanding of the hydrodynamic behavior in the
Figure 1. Simplified cross-flow naphtha-reforming (half) reactor scheme. Yellow arrow shows the direction of gas flow in the beds (zone B). Blue line: the direction of solids flow in beds. Purple line: the radial reduced temperature profile. The left side: drawing of solid interbeds transfer device in zone A’s with indication where the catalyst samples were obtained.
force between the catalyst bed and the vertical wall at the downstream side of the gas. With an increase of the gas flow rate, the frictional force may become large enough to impede distribution of the gas phase along the height of the catalyst bed, which might have a significant effect on the operating efficiency of a radial flow, as discussed in other analyses.21−24 On one hand, the poor axial distribution may influence global reaction conversion and selectivity, the temperature profile, and the catalyst activity profile, and hence, affect the quality and yield of product. Iranshahi et al.26 included the catalyst deactivation in the simulation. Continuous catalyst regeneration was modeled by Lee et al.,27 who reported good agreement with the data of the plant; these authors disregarded isomerization reactions and the solid fluid dynamics in the reactor. Stijepovic et al.28 proposed an improved bidimensional reactor model for a CFCCR unit; Hou et al.29 simulated the entire process in Aspen, and Lid and Skogestad30 determined the optimal operating conditions for this type of unit utilizing a Smith-type reforming model. Several of these studies simulated the reforming reactions and the fluid flow in a radial reactor using a simple catalyst deactivation function. In one-dimensional modeling, it was assumed that the gas stream passes through the particle bed as an ideal onedimensional radial flow; this is not the case, however, because gas flow is not ideal and there are axial gas flow components and end-effects in the catalyst bed.26 Poor distribution of the gas phase along the height of the catalyst bed has a large effect on the operating efficiency, because it influences activity, selectivity, temperature profile, and catalyst stability. Previous 6939
dx.doi.org/10.1021/ef300996z | Energy Fuels 2012, 26, 6938−6959
Energy & Fuels
Article
Figure 2. Sequence of spent catalyst mixing and deactivation (see Table 7 (given later in this paper) for yields). Data from a Commercial Plant. Four test runs were performed during operation of the unit to control the catalyst circulation effects: (1) deep naphtha and product analyses, (2) full operational data reconciliation, (3) instrument verification, (4) mass and heat balances were performed.39 Information about the mechanical design of the internal devices in the commercial reactors is not available. A large volume of representative catalyst samples were obtained from commercial plant during operation and during catalyst download at the end of cycle. These representative catalyst samples were obtained under the supervision of the author from specific locations in the reactors to avoid contamination; they were sealed and kept under inert atmosphere for further characterization. Data from the Microplant. The activity and selectivity of regenerated and spent catalysts were measured with naphtha in a microreactor laboratory plant. In addition, the spent catalysts were deactivated in the microreactor plant to develop data for adjusting the kinetic and deactivation constants, the “equivalent deactivation” procedure used was developed in previous studies33 and it reproduce the catalyst performance of an adiabatic fixed-bed reactor in an isothermal one. Catalysts deactivated in the pilot plant are designated as Spent*. To verify the catalyst mixing effect a small-scale apparatus (cold model) was developed and used for partially mixing the solid particles picked up from commercial plant before uploading the microplant reactor by layers. (See detail in the flow scheme in Figure 2.) Deactivation was performed according to the following protocol: 50 g of regenerated catalyst (Reg) from the commercial plant was uploaded in a small fixed bed (axial flow) isothermal reactor and heated to the reaction temperature under hydrogen stream containing 2 ppm of hydrogen sulfide and 20 ppm of chlorine. The catalyst activity at the start of the run was tested using commercial naphtha and hydrogen at three temperatures (510, 520, and 530 °C), and two space velocities (1 and 1.2 h−1); the other operating conditions were 1 MPa of pressure, 3.5 H2/HC ratio (molar). Then, Reg catalyst was deactivated at 510 °C by treating commercial naphtha and pure hydrogen for 9 days at 5.7 h−1, 0.6 MPa, and 2.5 H2/HC. At the end of deactivation period (EOR), the activity was tested again at the same three temperatures and two space velocities mentioned above. The product of this first reactor deactivation cycle is intermediate reformatted naphtha, designated (P1). The catalyst was cooled down in hydrogen at 35 °C, the reactor was opened under inert atmosphere, and the catalyst particles were downloaded from the bottom.
transfer zone is important for computer-aided modeling and for optimum design of internals. This paper develops a two-zone mathematical model to simulate the behavior of continuous radial stacked-bed reactors for the reforming of naphtha. This two-zone model takes into account fluid-dynamics effects. For the reactions, it uses a lumptype kinetic model, and, for the deactivation of catalyst, it employs two functions of the active site activity. Information about kinetic and catalyst deactivation was developed in a pilotplant test; in addition, commercial operational data and some of the information developed in a cold model are presented to explain the simulation results and the effect of variables.
2. EXPERIMENTAL SECTION 2.1. Commercial and Microplant Data. A commercial stackedbed continuous reforming unit was operated for 18 months before turnaround. The four-bed unit contained 140 m3 of Pt-type commercial catalyst and processed 150 m3/h of naphthenic (C6− C11) heavy naphtha to produce reformate with an octane number between 95 and 102, depending on the gasoline pool requirements. The feed contained ∼8% aromatics, ∼32% naphthenes, and ∼50% paraffins. The general description of the reactor system can be seen in the open literature.1 A simplified scheme of the reactor system is shown in Figure.1, along with an indication of the point where the catalyst samples were obtained. The flow rate of catalyst (circulation) was set up to average between 950 kg/h and 1400 kg/h; nevertheless, transfer was sufficiently continuous to prevent regeneration of the catalyst by batch and to seal each of the vessels from the interbed direct gas flow; periodically, fresh catalyst was added to compensate for catalyst attrition and similar amount of fine particles withdrawn from the regeneration system. In this study, researchers at times obtained and characterized samples of regenerated catalyst (Reg) returning to the inlet of the upper reactor samples of spent catalyst (Spent 4c) leaving the fourth reactor (but ahead of routing to a regeneration stage) to determine the level of activity of the solid and the amount of fines formed by attrition. At the end of the run, the plant was shut down for a normal turndown. The catalyst was downloaded, and a 1kg sample of four spent catalysts that might represent the top of the bed catalyst in each of the reactors bed was obtained (Spent 1a, 2a, 3a, and 4a). A special procedure developed in-house39 for downloading the catalyst was used. In addition, seven radial samples (1 kg each) were taken manually from the bottom of the reactor near the transfer line in each of the four beds (Spent 1b1−7, 2b1−7 3b1−7, and 4b1−7; the subscript “1” denotes the external layer of catalyst and the subscript “7” represents the internal one). 6940
dx.doi.org/10.1021/ef300996z | Energy Fuels 2012, 26, 6938−6959
Energy & Fuels
Article
obtained from a commercial reactor mixed and deactivated in a pilot plant. In addition, physicochemical information about spent and deactivated catalysts is obtained. Liquid products were analyzed using ASTM-P-0031-FAP (PIONA)40 technique, and gas products, using a standard gas chromatography (GC) method with a Si22 capillary column. Identification of 92 major hydrocarbons in naphtha and reformate was performed with a Hewlett-Packard Model HP 5970B MassSelective apparatus. The ion source of the mass spectroscopy (MS) was operated in electron impact (EI) mode with the electron energy at 70 eV and using the electron multiplier potential 1800 Detector. The split injection (1 μL naphtha or reformate) ratio was 1:50, and the injector temperature was 250 °C, with helium as the carrier gas. The column temperature was maintained at 40 °C for 8 min after injection and then programmed to maintain 200 °C for 10 min. Elemental analysis for the solid was performed by microcarbon and atomic adsorption (S, Cl) analysis. In addition, the commercial reformate was analyzed using a precolumn of molecular sieves (13X) and a back-flux system to feed conventional PIONA analyses. The precolumn produced a rough separation of C5−C6 and then the back flux sent the C7+ into a PIONA-GC-MS analyzer. The other precolumn separated C5−C8, sending the C9+ into the PIONA-GC-MS analyzer. The C5−C6 precolumn was operated with a temperature ramp-up between 40 to 100 °C, and the C5−C8 with a temperature ramp-up of 40 to 150 °C. The total analysis of each sample took 5 h. The detailed analysis provides a better resolution of C9+ isomers and alkyl aromatics for octane number calculation; the analyses were processed through a data-reconciliation procedure to reproduce the PIONA-MS analyses of the reformate. In particular, carbon balance was checked to ensure that the two analyses were consistent and that they reproduced the boiling point curve. Reformate from test runs #1 and #2 was also extracted using an ethylene glycol solvent (2/1 ratio of hydrocarbon (HC) to solvent) at 30 °C, contacted for 2 h in a autoclave, settled for 1 h, and separated into two fractions. The extract and the refined streams were washed with water four times, separated, dried by molecular sieve, and redistilled. PIONA analyses were performed in the two phases. Mass balance indicated only 1.5%−2% losses of light hydrocarbons. 2.2. Spent Commercial and Micropilot Plant Deactivated Catalyst Characterization. Four average samples of catalyst (5 kg each) were obtained from the inlet (Spent 1a, 2a, 3a and 4a), and 28 small (radial) samples (1 kg) from the outlet to the inlet of the beds (Spent 1b1−7, 2b1−7, 3b1−7, and 4b1−7) in commercial reactors; they were obtained during the shutdown and catalyst downloading of the unit after a normal operation cycle. The proportional blend of seven radial catalyst samples that reproduced the radial amount of solid in the reactor is called the average sample (Spent 1bav, 2bav 3bav, and 4bav). The amount of coke on 20 catalyst particles from each of the radial samples (1−7)statistically selected from pilot-plant deactivation test (Spent*) and commercial plant (Spent)were characterized using a microcarbon analyzer. In addition, average S and Cl elements were determined usign an atomic adsorption technique in the same particles. Physical properties were measured on average samples by the nitrogen adsorption−desorption technique (surface, pore volume, and pore diameter). In particular, benzene/N2 effective diffusivity was determined using a Wick−Kallenbach-type cell.41 The results are reported in Table 1 and Table 2. To verify the effect of catalyst mixing on activities, the Spent 4b and Spent 4c samples were tested with naphtha and hydrogen as feed under high-severity operating conditions (520 °C, 1.0 h−1 space velocity (LHSV), and 0.9 MPa of pressure) to obtain product of around 100 octane. The product was analyzed by carbon number and reported in Table 3. 2.3. Solid Dispersion Measurement in a Cold Model. Solid transfer between reactors was simulated using a cold model assembly (Figure 3) formed by a mini-screw conveyor to feed solid from the top, perforated 3-m-tall concentric slice of a cylinder (0.3 and 0.14 open area) with gas cross-flow, a hopper zone divided by seven removable rows, an rotating valve to control the solid flow, short transfer line, discharging inverted funnel, and a short-ended cylinder
The seven axial portions obtained at end of run were designated Spent*1b1−7. The solid samples were characterized and then tested with naphtha to compare the product with those obtained with Spent 1b1−7 samples. The seven samples of catalyst obtained from the bottom of the R1 commercial reactor (Spent 1b1−7) were installed in the hopper of a laboratory-scale cold model transfer device. There, the solids were partially mixed using the procedure described in the next section. Of the resulting mixed solids, 35 g were uploaded by layers into microreactor 2, heated at reaction temperature (510 °C) under hydrogen stream (with H2S and Cl2 content) for 4 h. The catalyst was then tested with naphtha at three temperatures (510, 520, and 530 °C), and two space velocities (1 and 1.2 h−1), as mentioned previously. The catalyst was then deactivated by treating the product P1, collected in previous run (first cycle of deactivation), at 510 °C and 4 h−1 (a similar liquid space velocity to that used in the second bed of the commercial reactor) for 9 days in continuous operation. The other operating conditions were 0.6 MPa of pressure and 2.5 H2/HC ratio (molar). The products collected during this period on stream are another intermediate reformate, designated P2. At the end of the run (EOR), the catalyst was tested again with naphtha at the same three temperatures and two space velocities. The reactor was then cooled in hydrogen, and, at 35 °C, the reactor was opened under inert atmosphere, and catalyst particles were downloaded from the bottom in seven axial portions, designated Spent*2b1−7 (recall that the asterisk (*) indicates that the catalyst was deactivated in the pilot-plant sample). These samples were characterized for comparison with the Spent 2b1−7 samples. The seven samples of catalyst obtained from the bottom of the R2 reactor (Spent 2b1−7) were mixed using the same protocol described below. After mixing, 25 g of these solids were uploaded in a microreactor, heated, activated, and tested with naphtha and hydrogen at the three temperatures (510, 520, and 530 °C) and two space velocities (1 and 1.2 h−1) previously used. The catalyst then was deactivated by treating the product P2, collected in previous runs (second deactivation tests), at 510 °C for 9 days at a liquid space velocity of 3 h−1, pressure of 0.6 MPa, and 2.5 H2/HC ratio (molar). These operating conditions are similar to those of third bed of the commercial reactor. The collected product during this period on stream is another partial reformate called P3. At the end of the period, the catalyst was tested again with commercial naphtha at the abovementioned three temperatures and two space velocities. The catalyst was downloaded from the reactor under inert atmosphere in seven samples (Spent*3b1−7). These solids were characterized, and the results compared to those obtained with Spent 3b1−7 samples. The deactivation of Spent 3b1−7 samples was not carried out, because of contamination problems on the P3 sample in the micropilot plant. In summary, the procedure for deactivation and mixing of commercial samples in co-current operation (see Figure 2) is:
Test at start of run (SOR) Reg catalyst activities with naphtha at three temperatures → Deactivation of Reg catalyst with naphtha at 510 °C→ At end of run (EOR), test deactivated catalyst activities with naphtha → Recover Spent* 1b1−7 and P1 samples → Partial mixing of samples of catalyst Spent 1b1−7 → Test at SOR the mixed catalyst activities with naphtha → Deactivation of Spent 1b catalyst with P1 feed at 510 °C→ At EOR, test deactivated catalyst activities with naphtha → Recover Spent* 1b1−7 and P2 samples → Partial mixing samples of Spent 2b1−7 → Test the mixed catalyst activities with naphtha → Deactivation of Spent 2b with P2 feed at 510 °C. At EOR, test activities with naphtha → Recover Spent*21−7 and P3 samples → Partial mixing of Spent 3b1−7 catalyst sample → At SOR, test mixed catalyst activities with naphtha. This protocol produces solid and liquid samples that allow us to calculate the conversion and selectivity at the SOR (equivalent to the inlet of the commercial reactor) and EOR (equivalent to the outlet of the commercial reactor); notice again that original catalysts were 6941
dx.doi.org/10.1021/ef300996z | Energy Fuels 2012, 26, 6938−6959
Energy & Fuels
Article
operation are reported to explain the blending effect. Several mechanical designs (angles) and times at rest of catalysts in beds (the period of time for the operation of the valve) were used for the study of solids mixing at different throughputs. The mechanical design is based on the ideas mentioned in ref 16 and previous work in movingbed cold model studies.42 This information was employed to build a mini-mixing device that was used in pilot-plant deactivation studies. The paper discussing the cold model study is under review by another journal.31 The nickel-containing particles, impregnated with different concentration (by weight) of tracer (αNi), were installed between seven removable radial fences (rows) on the top of the upper cone (zone B1 hopper, Figure 3). The fences then were removed to allow the solid to move downward freely and the upper cylinder was installed; the upper cylinder was filled with a commercial catalyst (on top of the traced solid) and the gas cross-flow gas circulation was set up; hence, the solid was discharged at a constant head pressure through the upper cone, interconnection pipe and lower inverted cone into the lower bin where there is a pile of catalyst (without tracer) formed in the previous discharge. The nickel-containing particles fall into the top of the pile of solid and move to the side; in the lower cylinder B2 (bin), there are seven rows to separate the solid by layers at end of the discharges (see Figure 3). The radial profile of Ni was measured by chemical analysis of 20 particles present in each of these seven rows. During discharge, a constant level of solids was maintained in the upper hopper to operate in the so-called “mass-flow regime” at a constant head pressure. The solid discharge rate was explored using cone angles in the hopper between 35° to 38°; the acrylic glass wall− solid angle was 17° (measured using a Jeniker’s tester), and the open and close time for the rotating valve was 1−20 min. Solid flow rates were varied from 100 to 300 kg m−2 h−1. Although these rates do not correspond to the mass flow rate per unit of area of commercial reactors, they seem to reproduce the mixing effects, as will be discussed below. Notice that, in the commercial reactor and in the cold model device, the number of catalyst particles per row and the weight of solid decrease with the radius. The solid volumetric distribution in the cold flow model corresponds to those of commercial reactors. Therefore, the average value of the nickel concentration is calculated given the number of particles and the concentration of the tracer in each of the rows. The inlet average nickel concentration was adjusted to the value of coke in spent catalysts. Remember that deactivation of the Reg catalyst along the first commercial reactor (R1) generated a radial profile of coke on catalyst at the outlet (Spent 1b1−7); catalyst transfer to the second reactor produces a different coke distribution a the top of the second reactor (Spent 2a). Then, deactivation through the second reactor generates another radial profile of coke at the outlet of the second reactor (Spent 2b1−7), with coke always being higher in the external layer (inlet of gases) and lower at the axis layer (zone of gas recollection). The same happened with the radial profile of coke through the other reactors. The operating conditions of the transfer device were adjusted to reproduce the profile of coke at the top of the commercial reactors from R2 to R4 empirically. The mixing effect of Ai zones was simulated using a mathematical model based on a series of continuous stirred-tank reactor (N-CSTR) equations. These mixing
Table 1. Carbon Distribution per Families (PNA) of Compounds (510 °C, LHSV = 1 h−1) carbon/100 mol naphtha C6 molar distribution paraffins naphthens aromatics molar distribution paraffins naphthens aromatics molar distribution paraffins naphthens aromatics molar distribution paraffins naphthens aromatics molar distribution paraffins naphthens aromatics
C7
C8
Feed 16 28.5 7 9.5 6 15 3 4 ProductSpent 1bav 29.91 33.84 14.22 16.5 3.09 1.24 12.6 16.1 ProductSpent 2bav 27.91 33.65 14.5 15.75 4.01 1.8 9.4 16.1 ProductSpent 3bav 24.7 32.8 13 15.5 4 2 7.7 15.3 ProductSpent 4bav 23.1 31.83 12 14.4 3.1 2.13 8 15.3
27.2 9 13.2 5 Catalyst 23.25 6.2 2.8 14.25 Catalyst 23.9 7.1 1.4 15.4 Catalyst 26.2 8.3 1.5 16.4 Catalyst 27.8 9.8 1.8 16.2
C9
C10
14.9 7 3.9 4
8.2 5 2.2 1
10.18 0.82 2.02 7.34
2.782 0.122 1.26 1.4
11.3 1 2.3 8
3.24 0.44 1.15 1.65
12.06 1.66 1.18 9.22
4.26 1.23 1.23 1.8
12.49 1.8 1.35 9.34
4.8 1.3 1.3 2.2
with removable rows to collect the solid. The apparatus also has solid sampling ports, a mass (solid) flowmeter, gas metering, and gas and solid sampling. The solid is discharged at specific rates controlled by the rotation of the valve during a pre-established time; then, valve operation is stopped and the bed is settled during a certain time. The solid discharge was done many times to evaluate the reproducibility of the operation. The mixing effect was evaluated using Ni-acetonateimpregnated particles as the tracer to represent the coke content in spent samples; the particles have the same density, shape, and particle diameter as the commercial catalyst. The repos angle of the solid and the Ni distribution on the solids stored in the lower bin were measured under different operating conditions (solid rate circulation, time at rest, and gas flow rates). The result of this cold model simulation was used to develop a solid flow circulation model based on Eulerian equations,31 in which the different phases are treated mathematically as interpenetrating continua. In particular, the author determined the effect of time at rest in the solid “pining” bed density and pressure drops. Here, the solid mixing results that allowed reproduction of the partial mixing and the time at rest for the catalysts in commercial
Table 2. Spent and Deactivated Catalysts in Fixed-Bed Operationa
a
catalyst
Reg
Spent1bav
Spent2bav
Spent3bav
Spent4bav
Spent 4c
coke (wt %) sulfur (wt %) diffusivity (× 102 cm2/s) surface area (m2/g) pore volume (cm3/g) aromatics (wt %) naphthenes (wt %) paraffins (wt %) octane MON
