Solid Catalyst Alkylation of C2âC3 Olefins with ... - ACS Publications

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Solid catalyst alkylation of C2-C3 olefins with isobutane in the presence of hydrogen using a slurry transport reactorhydrocyclone-regenerator system and PtSO4TiZr/SiO2 catalyst. I. Alkylation in Pilot plant and Simulation of STR-HCS system Roberto E. Galiasso Tailleur, Sergio Rodriguez Dominguez, Carlos E Farina, and Sylvana Derjani-Bayeh Energy Fuels, Just Accepted Manuscript • DOI: 10.1021/acs.energyfuels.7b03015 • Publication Date (Web): 26 Dec 2017 Downloaded from http://pubs.acs.org on December 27, 2017

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Solid catalyst alkylation of C2-C3 olefins with isobutane in presence of hydrogen using a slurry transport reactor-hydrocyclone-regenerator system and PtSO4TiZr/SiO2 catalyst. I. Alkylation in continuous Pilot plant operation and Simulation of STR-HCS system. Roberto Galiasso Tailleur*, **1, Sergio Rodríguez*, Carlos Farina** and Sylvana DerjaniBayeh* *Grupo TADiP, Departamento de Termodinámica y Fenómenos de Transferencia, Universidad Simón Bolívar, AP 89000, Caracas 1080, Venezuela,** HyPro Consultant 4250 Corrine Dr., Suite 204, Orlando, Florida USA E-mail: [email protected] ABSTRACT A continuous alkylation plant was simulated to evaluate the capability of a newly designed acid solid catalyst PtSO4TiZr/SiO2 to convert light olefins (C2= and C3=) and isobutane into alkylate in presence of hydrogen. This part of the process consists in a three-phase slurry transport reactor (STR), a hydrocyclone-settler (HCS) operating in series, with recycle of unconverted reactant and regenerated and fresh catalysts. Data for alkylation were obtained in batch reactor and in pilot plant tests at different gas flow rates, temperatures, pressures and catalyst particle sizes. Kinetic and deactivation rates and fluid dynamic data were obtained using the pilot plant continuous operation. Fresh, spent and regenerated catalyst were characterized using different techniques to explain its selectivity and deactivation. Commercial size plant was design and the simulated to determine the impacts STR operating variables in the alkylate cost. The study determined the kinetic and deactivation rates parameter for main reactions, as well as the effect of operating variables and particle size in the performance of the catalysts. Simulation provide information to discuss the behavior of STR reactor, and the impacts of the operating variables in the cost of alkylate. The rates of apparent in-series alkylate production depend on order one in intermediates and olefins concentration, catalyst activity and rate of diffusion in mesopores; the rates of soluble coke production depend on order one in intermediaries and 1.5 in olefins concentration and is inversely proportional to hydrogen partial pressure. Kinetic and deactivation model

is

independent of the reactor used. Catalyst aged with the number of cycles alkylation-

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regeneration; its activity is inversely proportional to the soluble coke content. Inlet temperature, olefins-to-isobutene mole ratio, hydrogen-to-olefins mole ratio, catalyst make-up are the main operational variables that impact the cost of alkylate. Sensitivity to kinetic rate parameters, fluid dynamic model, and particle size are analyzed. There is an optimal coke built up in solid alkylation that minimize the alkylate cost.

Keywords: Solid alkylation, slurry-transport reactor, hydrocyclone-settler, light olefins alkylate

1. INTRODUCTION The production of a “green” gasoline using solid alkylation is one of the most challenging tasks of process developments. Current commercial production of alkylation of olefins is accomplished using strong liquid acid (HF and H2SO4), these technologies have good selectivity and activity but generates toxic sludge (H2SO4) or have strong safety limitations (HF). A competitive solid catalyst should be active, stable, and regenerable. The Jong et al (1996)1 among others mentioned that a well-stirred slurry reactor is the prefer technology that can extend catalyst lifetimes in solid alkylation of isobutane with 2-butene. Low olefin concentration helps to control the formation of carbonaceous deposits, the main drawback of the technology. The solid alkylation catalysts and reactions of isobutane with butenes were review by Corma (1996)2 and Lercher (2004)3, among many others. Weitkamp et al. (1987)4, Pater et al. (1995)5, Flego et al. (1995)6 and Feller et al. (2003)7 have discussed the catalyst deactivation in different solid and they conclude that coke formed is composed mainly of high-molecular-weight and highly branched paraffinic molecules/ oligomers/ polymers with one or several double bonds. In addition, there is several fundamental kinetic studies performed in continuous stirred tank reactor (CSTR), among those is important to mention the work of Nivarthy et al. (1996)8 and Martinis and Froment (2008)9. The later authors summarized in their Fig. 5, eight fundamental mechanistic steps for solid acid alkylation of butenes on zeolitic catalyst at moderate temperatures. These elementary steps are the alkene protonation, alkene deprotonation, intramolecular hydride shift, methyl shift, protonated cyclopropane (PCP) branching, oligomerization, β-scission, and intermolecular hydride transfer. The reactions network for this process is the result of many elementary steps


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controlled by carbenium ion chemistry. Cracking shifts the carbon number distribution toward shorter chain lengths. On PtZrTiSOx/SiO2 catalyst10,11 the alkylation mechanism of light olefins is different than in other zeolites and it occurs through a simple series of olefin addition to an carbenium ion follow by dimerization ciclization and dehydrogenation. Gonzalez et al. (2006)

10

Galiasso et al. (2008)11, Galiasso (2008)12 developed a process

scheme to alkylate butenes with isobutane in presence of hydrogen; them used a slurry transport reactor (STR), a hydrocyclone gas-solid-liquid separations and stripping stage (HCS) and a catalyst regeneration stage13 to perform the alkylation (see Fig. 1, left side.) The reaction zones are couple to conventional separation heat exchangers and recycling stages; unconverted olefins hydrogen and isobutane are recovered, recompressed, heated and recycled to the STR (Fig. 1, right side.) Simulation of the potential commercial size process is based on thermodynamic properties, fluid dynamics and mass and heat transfer correlations, and apparent kinetics and deactivation rates model. The latter model was first generated in a laboratory batch reactor and then verified in a continuous CSTR pilot plant; a PtZrTiSOx/SiO2 was used as catalyst, operating under hydrogen. There is no information available in the literature about the catalyst, the mechanism of alkylation and deactivation to compare with. A simulation program obtained the rate parameters and verified the mechanism. The effect of variables were used to discuss the simulation results. The STR reactor is like those employed for coal liquation and more recently by Fischer Tropsch synthesis, but with internal recycle (spouted bed.) Slurry type reactors were discussed early in time by Parulekar and Shah (1982)14 and more recently by Sie and Krishna (2000)15. Galiasso and Andretti (2008)16 adapted this reactor technology for butenesalkylation and here is extended to ethylene and propylene alkylation with important modification. The simulation of a commercial-size STR-HCS equips were performed using the new apparent kinetic and deactivation rates expressions, effects of mass transfer in pores and


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previous information about fluid dynamic model16. Another “ad hoc” reactor program was developed to simulate the slurry transport reactor-hydrocyclone stripper system (STR-HCS, integrated to a Regeneration reactor (Fig. 1, left side) and a Fractionation stages (Fig 1 right side). The simulation program provides information about product yields, profile of temperature, catalyst activity, effect of particle size and coke deposition in the alkylation reactor, as well as their impact in the cost of alkylate. The effect of correlation used are discussed. The results of the STR-HC simulation at different operating conditions were analyzed based on alkylate cost. The second part of the paper describe the pilot plant experiments and simulation of the fluidized bed reactor (FBR) used for catalyst regeneration13. EXPERIMENTS AND SIMULATION 2.1 Experiments The following sets of experiments develop the data needed for the pilot plant operation and commercial size reactor simulations. 2.1.1 Catalyst preparation for pilot plant tests. Briefly, PtSOxZrTi/SiO2 catalyst was prepared by a gel co-precipitation of beta-diketone (2,4pentane-dione) tetra-ethyl-orthotitanate, and a tetra-ethyl-orthozirconate (Aldrich) in ethanol. Then N-methyl-2-pyrrolidone (NMP) and a large molecular weight Si hydrogel were added under intense stirring. The precipitate was washed with ammoniac solution at 298 K; then it was dried under vacuum at 350 K for 8 hours to form hard spheres with a particle size between 40-60 microns. The solid was treated in oleum for 4 hours at 313 K, filtered, washed with methanol, dried, and impregnated with diamine platinum dichloride aqueous solution. Finally, the solid was dried at 393 K and calcined in air at 523 K for 12 hours and activates at 545 K for another 12 hours. Fresh (D) catalyst were used in the alkylation pilot plant that generated D1 spent catalyst; then the latter was regenerated in another pilot plan that produce D11 catalyst. 95% of latter plus 5 % of fresh catalysts are treated again in the alkylation plant to generate D12 sample. The process continues until obtaining 45 kg of D33, a catalyst with three cycles alkylation-regeneration. After several cycles though alkylation – regeneration stages, the catalyst is progressively deactivated by accumulation of CS2-insoluble coke in micropores17. Fresh, spent and regenerated catalysts were characterized. Carbon and acid sites


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content, adsorption and diffusion of reactant and product and physical properties of spent (D33) and fresh plus hydrogen-regenerated catalyst (D43,) are reported in Table 1. 2.1.2 Alkylation reaction: initial activity and long-term test. Kinetic and deactivation study. The apparent alkylation kinetic and deactivation rates were study in a small batch reactor (Fig. 2a, batch in solid) at different contact times in successive tests (Runs) using D33 catalyst; different temperatures (353, 363 and 373 K), pressures (0.8, 1.4 and 1.8 MPa), average catalyst particle diameters (50, 100 and 150 microns) were used. Combinatorial ratios of olefins / isobutane (O/IB: 0.2, 0.3 or 0.4), ethylene / propylene (C2/C3: 0.3, 0.5 or 0.8) and catalyst /hydrocarbons (Cat/HC: 0.04, 0.08 or 0.12) employed. The effects of products of reactions were determined by adding at t=0 one of products together with iC4, to have a ratio of either 0.1, 0.2 or 0.3. The ratios studied were DMB (dimethyl-butane) / iC4 (isobutane), TMB (trimethyl-butane) / iC4 and HCH (hexyl-cyclohexane). The hydrocarbons (iC4, olefins, hydrogen and any of products) are fed into the reactor at the operating temperature and pressure and then, solid were added under smooth stirring (100 rpm). After 0.1 hour of contact time (Run-1), the catalyst, gases and hydrocarbons were downloaded and cooled to 300 K for analyses. After that, other tests were performed at the same operating conditions but at either t=0.2h (Run-2), t=0.3 h (Run-3) or t=0.45 h (Run-4) of contact times. Gases, liquid and solid samples were analyzed by GC-MS and micro-combustion respectively to determine olefins conversion, selectivity and coke formed on catalyst after each of runs. Fig. 3a (points) reports the information about the effects of residence time and temperature in olefin conversion and coke production. Fig. 3b (bars) shows, as an example, the effects of particle size diameter, pressure, temperature and products of reaction in olefins conversion and coke built up, all of them at constant contact time (t=0.3 h); all other operating conditions are indicated in the caption. The performances of the batch reactor were simulated by guessing kinetic and deactivation rates constants and by solving mass balance equations depicted in annex I for a isothermal reactor, at different operating conditions. The simulation program adjusts the constants, using a Genetic algorithm (GA) program to fit the experimental data. Activity of catalyst was calculated by the model at t=0 for catalyst tested with different initial coke content, catalyst particle size and one of products of reactions. Table 1 reports, as an example, the initial activity of spent catalyst D33 and the make-up (D34)


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formed by 5% fresh plus 95% regenerated catalysts. Activity was always measured by extrapolating the data of conversion as a function of contact time to t=0 in the batch reactor; the activity tests were performed at 373 K, 1.4 MPa, and with molar ratios of 1 for C2/C3 and 1 for O/iC4; in addition, a PH2 of 0.72 MPa, dp of 50 microns and Cat/HC of 0.08 wt. were used. Fluid dynamic of the pilot plant reactor. To develop an apparent kinetic and deactivation rate model, independent of the reactor fluid dynamic, several tests were performed. Bubble size distribution (BSD) in a five-liters isothermal CSTR operating in steady state conditions (pilot plant, Fig. 2b) were determined in 22 blank tests using catalyst in hexane plus H2, without reaction; for that either 4, 8, 12 or 20 % of solid content, 50, 100 or 150-microns average particles sizes were employed. Slurry was stirred by convex two-pitched blade turbines (top and bottom, see Fig 2b). One perforated gas-sparger with 20 holes distributed hydrogen into slurry phase. The effects of gas flow rate and slurry rate of stirring were determined in 15 tests, operating at different levels of conversions in the CSTR and using base case operating condition. Slurry temperature and total pressure were controlled at ±0.1 K and ±0.01 MPa respectively. Inlet-outlet delta of pressure was continuously recorded. A DANTEC 62 X modular series laser Doppler anemometer, operated in a forward scattermode, was used to determine the effect of bubble size, type of impeller and rate of stirring the slurry on olefins conversions. Data validation and signal processing were carried out using Sunteg 629 Doppler signal-processor and Honeywell 423 DPC. Gas density in the freeboard was measured by anemometry and solid content in the gas leaving the reactor by weighting the filter bags. Olefins alkylation in long term tests. Long-term tests were performed in the CSTR at three different temperatures: 375, 380 and 385 K, three pressures: 0.8 1.4 and 1.8 MPa, using either 50, 100 or 150-microns average particle size catalysts; 95% of D33 plus 5% of D catalyst were used as catalyst make up. Solid was automatically added and withdraw each minute using an automatic (computer-drive) program that control two hightemperature and pressure valves. Pressure, delta of pressure and temperature were recorded as a function of time in stream; variation of local pressures and bubble sizes was determined with a 1 mm A231thermo-sensor probe and a three-tips laser probe respectively. Reactants have the following molar


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proportions used as pivotal values: C2/C3: 1, O/IC4: 0.3 and (H2/O): 5 (molar ratios). Gas (Tg), liquid (TL) and solid (TS) residence were adjusted to keep conversion around 40-60% at all operating conditions. Catalyst samples, withdraw from the pilot plant, were tested daily in batch reactor to determine their initial (t=0) activity and to measure products yields at the same operating conditions. Daily activity of catalyst had been used to follow catalyst deactivation and predict the amount of catalyst to be added and withdraw from the pilot plant, during the long-term tests, to keep the activity at certain level (between 40 to 60 % conversion). Catalyst activity as a function of coke content were measured each two-hour at steady state and during transient mode of operation. The value reported in Fig. 4 is a daily average; left Yaxis shows the results of olefins conversion to alkylate, and in right Y-axis the conversion of olefins to “coke” plus paraffins, both plotted as a function of time on stream. When the particle size was changed after 32 days in continuous operation, the catalyst inventory was downloaded completely, and new catalyst added with different particle diameter. Then, the plant was operated in a transient mode until the new catalyst inventory and conversion were achieved (three days). By simplicity the new particle size was tested at three operating conditions. After that the procedure were repeated for another particle size. The catalyst inventories that had been withdraw were divided in 280 sets of five particles and coke analyzed. The plot of coke as a function of number of particles were used to determine the “age” of catalyst in the CSTR reactor. Samples of spent catalyst were accumulated, characterize and then used to perform the regeneration studies13. Olefins conversion, alkylate, gas and coke production were determined by analyses of gases liquids and solids withdrawn at different time on stream. GC analyses of gas and liquids were performed on line using a GC-FID (HP 6890, PIONA column). Coke was measured by microcombustion and

13

CNMR analyses, acid sites determine by pyridine desorption (Thermo-

balance and FTIR analyses17), soluble-coke measured by CS2 extraction and physical properties determined by XRD, TEM, nitrogen (BET), benzene, iC4, DMB and TMB adsorption-desorption; Description of methods used in the characterization of spent catalysts can be seen in previous contributions10,12,17.


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A kinetic and deactivation rate model, were proposed using a set of parallel-series apparent reactions (lumps, equation 1-8.) Previous kinetic and deactivation rate studies17 performed with ethylene and propylene have shown that the rate of reaction depend on olefins and adsorbed isobutene [iC4]+ or intermediates (isoparaffinic species [iC6 -C9]+) concentrations; the stoichiometry and heat of reaction of alkylation are shown by (Eqs. (2) and (3)). Heavy (isoparaffinic) intermediates ([iC6 -C9]iP+) are hydrogenated and desorb as light alkylate, or form adsorbed-cycloparaffins ([iC6 -C9]N++) at the meso-micropores intersection17,18, or dimerize to form another adsorbed heavy intermediates ( [C12.5 H 25 ]iP+ + and [C12.5 H 25 ]+N+ ) that hardly desorb, leading to catalyst deactivation. Most of monomers and dimers present in micropores were unable to desorb and they are dehydrogenated, cyclized and polymerized in situ to form high molecular weight aromatics { [C12.5 H 25 ]+Ar+ }m polymers. Heavy adsorbed compounds present in mesopores are mention here as soluble-coke (coke that could be extracted by CS2 and be regenerated) and those that are placed in micropores as insolublecoke. The previously found mechanism of deactivation (series parallel reactions17) was simplified here by grouping dimerization, dehydrogenation reactions that form coke into a simple expression easier to simulate (equation 4 and 5), and their apparent kinetic rates constants obtained. Hydrogenation of olefins form light paraffins (Eqs. (9-10)) and light alkylate, decreasing C2C3 olefin concentration, and controlling coke [C10-16] formation on the catalytic surface. Pt promote the hydrogenation via hydride addition of some adsorbed light intermediatescompounds (coke precursors, Eq. 11) forcing them to desorb.

Apparent kinetic rates model. For olefins conversion between 40% and 60% the model use following set of eleven parallel reactions and their heat of reactions: fast fast iC 4   → [iC +4 ]i + [ H − ] , H 2    → [H - ] + [ H − ]

k111

C6

k211

(1) C8-9

k5 0 (r1,r3,r6): C =2 + [ iC +4 ] i k 1 → [C 6+ ] i  k3,  → [C 8+ + C 9+ ] i ; ∆ H 273 C2=, C3=

k111

C7

k211

C9-10


= − 83 .7 kJ mol of C 2

(2)

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k2 4 ,k 6 (r2, r4, r7) C 3= + [ iC +4 ] i → [C 7+ ] i  k  → [ C 9+ + C 10+ ] i ;

∆H

C2=, C3=

0 273

(3)

= − 104 . 6 kJ mol C 3

;k slow ; k + + (r5): 2.9C 2= + [ C 8+ ]i + [ C 9+ ]i + [C10+ ]i  -H ; 0  → [C 10 -14 ] Nap    → [C 10 -14 ] Aro ∆H273 ≈ 34kJ 2

7

12

;k slow + + (r8): 2.9C 3= + [C8+ ]i + [ C 9+ ]i + [C10+ ]i H ; 0 →[C 11 → [C 11 -15 ] Nap   -15 ] Aro ∆H273 ≈ 34kJ 2

(r9):

8

mol H2

mol H2

(4) (5)

0 ; ∆H 273 = −50.3 kJ mol H 2

(6)

k10 0 (r10): C3= + H2  →C3 ; ∆H 273 = −50.3 kJ mol H2

(7)

k9 C =2 + H 2 → C2

k1,2,3,4

11 0 →Ci ; ∆H 273 (r11): [C6+-15]i +[H- ]  = 0 kJ mol iC 4

Secondary and tertiary re-alkylation reactions are between adsorbed intermediates [iC6−10]+ and olefins present in liquid or gas phase in mesopores and micropores; the rate of re-alkylation reactions compete with the rate of hydrogenation-desorption (k11) that form iC6−10 alkylate; these rates depend on molecular size of the adsorbed intermediates. Dimerization of iC4 and cracking of heavy compound reactions do not play important roles. Very small amount of alkylate are produced by using n-butane instead of isobutane as feed at the current operating conditions. Additional information about the mechanism is available in reference17. The previous adsorption-desorption studies18 of binary sorbents (Hex-IC4; Hex-TMB; Hex-IC4 and Hex-HCH) provided information about the partition coefficients that are included in apparent kinetic rates constants. The equations used to simulate the pseudo-isothermal semi-batch reactor are shown in annex I and the methodology described in ref. 17. The apparent kinetic rates expressions for this type of catalyst are reported in Table 2, and kinetic rates constants in Table 3. Fig. 5 shows the product distribution obtained at different residence times in pilot plant. The fluid dynamic measures indicate that the pilot plant can be simulated as a pseudo-isothermal CSTR. Apparent reaction rates constants for all reactions (ki, Ei), deactivation constant (a) and particle efficiency factor (η) were determined by using the batch and pilot plant reactors data; A genetic algorithm program, seeded with the kinetic rate parameters obtained in batch mode of operation, were used to adjust the above mentioned model exp parameters to the minimum of objective function ( (CCexp 2+C3 − CC 2+C3 )100/ CC 2+C3 ) and obtain the

best set for ki,, Ei, a and η values 9 (see below.) Fig. 6 a, b and c depict predicted conversions of olefins into alkylate, coke and paraffins against experimental data generated in


(8)

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the long-term tests. Fig. 7a shows the predicted delta of catalyst activity versus measured values at steady state, and Fig. 7b shows the same in transient mode of operation. 2.2 Process simulation 2.2.1 Process flow diagram (Fig. 8). Hydrogen, isobutene, olefins and solid catalysts were fed at the bottom of the riser in commercial size STR. Slurry and bubbles move upward through the riser and only the slurry come down through the downcomer (space between vertical bundles of tubes and the reactor wall.) There is random mixing in the riser produced by the ascending bubbles (hydrogen, light paraffin’s, and olefins) confined by the bundle of cooling tubes, which drags up the solid particles and liquid by the center of reactor and induces a movement of denser phase with a fountain type effect on top (like a spouted bed reactor.) Above there is a freeboard zone that recover liquid droplets and particles; after the disengaging zone, hydrogen, light paraffins, and olefins are cooled and separated from the solid the hydrocyclone-stripper and went to the fractionation; the slurry leaves the reactor by the side toward the first hydrocyclone settler (HCS) device. In the HCS the slurry phase moves downward in spires at the wall and the gas-liquid- fines particles disengage by the center due to centrifugal forces, delta of densities and additional effect of bubbles of hot hydrogen; hot gases, injected at the bottom, vaporized hydrocarbons, liquid and fines of solids move upward in spires at the center (vortex finder). At two-third of vessel high, gases disengage while liquid plus fines is laterally withdrawn and send to a second HCS device. There, liquid is again stripped from the solid by adding hot hydrogen at the bottom. Gas holdups are closely associated with the size, number and the frequency of bubbles that that depend on pressure, temperature and fluid properties. Fig.8 is a zoom of the reacting system showing details of the interconnection STR- HCSFBR. Hydrocarbon and hydrogen gases stream leaving at the top of the STR and at the first and second staged of separation (HCS) are cooled and separated. Gases, rich in hydrogen, are purged, recompressed and blend with hydrogen make-up before being recycled to the STR inlet. Solids are recovered from the bottoms of first and second HCS and fed into densephase of a fluidized bed reactor (FBR) to be regenerated in presence of high-temperature high-pressure hydrogen. Detail of this process is described in second part of this paper13. Regenerated catalyst is purged and blend with fresh catalyst before being recycled to the inlet


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of STR. Liquid product from the second separation stage plus condensed olefins from the top of STR are forwarded to fractionation stages (Fig. 1b) were 97.5% of isobutane, and 89% C2 and C3 unconverted hydrocarbons are recovered. 2.2.2 Material and energy balance for the STR. To simulate the three-phase STR system, a sets of fluid dynamic, gas-liquid equilibrium, and simplified mass and heat balances equations, together with some catalyst properties are used. Fig. 9 shows streams and nomenclature used for simulating the solid alkylation and separation stages. STR simulation model used rates of reactions expressions depicted in Table 3 ( Eqs. (9) to (19)) and correlations shown in Table 4. Conversions are defined by equations (21 to 31.) Relationship between the molar flow rates at inlet (Fi) and outlet (F′i) and the different conversions are given by the Eq. (32) to (47). The two sets of conversions (C2: X1, X3, X5 X7 and C3: X2, X4, X6, X8) are defined for ethylene 1, 3, 5, 7 and propylene 2, 4, 6, 8 reactions respectively. The overall flow rates F° is a function the inlet (Fi) and recycled (RF′i) molar flow rates (Eqs. 32 to 47); activity is proportional to delta coke (Eq. 48). CC2,L and CC2,s are reactant concentrations in liquid and in solid wall of mesopores, Cip is the reactant or product concentration in the liquid that fill the pores which is at equilibrium with the concentration of the adsorbed species i; partition coefficient (or equilibrium constants) are K oi , L − S for liquidsolid and Koi ,G−L for gas-liquid. Values of these constants for light alkylate were obtained by Rough and Galiasso17-19. Reactants and products must be transferred from gas to liquid and then to solid surface. These internal and external mass transfer hindrances are accounted by using the two-film theory. Simulation model assumes that at steady state, the rate of reactants migration to the interface equal the rate of reactants reaction in pores, or the rate of products going back to the liquid phase. The effectiveness factor, η, corrects the apparent kinetic rates expressions for differences in olefins concentration between the catalyst external surface and those inside the pores: 1 ηi = 3 / φi (1/ tanhφi − 1/ φi) (49); φ =

rp

ralk (50) C6−9 Deff f


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There is a consensus in solid alkylation that the slow rate of hydrogenation, desorption and products diffusion out pores form coke. Yoo et al. (1998)20 and Weitkamp, (1999)21, among many others, had discussed the problem of slow desorption of alkylated products in zeolites channels. The values of effective diffusivity Deff for olefins and C6-C10 used in equation (50) were experimentally calculated from data generated in the batch reactor using particles of either 50, 100 or 150 microns diameter. Then, ηi value where confirmed using pilot plant data. The effective diffusion coefficient in mesopores are compared with those values measured without reaction by Rough and Galiasso19 and those reported in the lierature22,23. Here it was assumed that the mass transfer resistances in gas phase, gas liquid-interphase and in liquid phase are negligible. Mass transfer ( k SL − S ) and heat transfer ( ζ SL− S ) rate constants at the interphase liquid-solid are obtained from Ramachandran and Chaudhari19 and they are similar for all the species. Hydrogen and hydrocarbon concentration and molar volume of liquid are calculated, point by point along the grid used for integrating the heat and mass balances equations, using Peng- Robinson EOS equations (see below). The mass and heat balances of the plug flow recycle reactor model (STR) are: Gas phase: Mass balance for C2 olefins dCC 2G hG kGL aGL PC 2 = ( − CC 2, L ) (51) dV r n g vo KC 2

Slurry Mass balance for C2 olefins: 1.5 0.3 dCC2 kSL−S (CC2,L − CC2,S ) aη1 (k 1CC2 CCIC4 + k3CC2 (C6 + C7 ) + k5aη1CC2 (CC8 + C9 ) + +ak9η1CC2CH 2 ) = + E [ A + KIC4CIC4 + KC6−10CC6−10 ].E dV r

(52)

+ k7CC1.52CC8−C10 /(1+ 4CH 2,l )w/ wo

Similar balances were used for the other families of hydrocarbons and hydrogen (C3, iC4, C6…C15, H2). There, the parameters are: B=

R C 00 + C 0 (1 + R1 + R2 )vo (53), E = R2 vo (54) R1 = q1 (55), CC1 2 = 2 C 2 C 2 (56), hL ( R2 + 1) vo hL

R2 =

f 1 q2 , (57) CC2 2 = R1CC 2 + CC 2 (1 + R1 ) (58), vo ( R1 + R2 + 1)

Solid phase


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Energy & Fuels

W2f = W5f (1 + R1 + R2 ) =

a2 =

V1hl ρ SL V hρ vo ρ L (1 + R1 + R2 ) (59); τ = (60), τ s 2 = 2 l SL (61) s1 1 v0 ρ L (1 + R1 + R2 ) v0 ρ L R2 ( − 1) Fc

(1 + R1e k d E (t )τ s1 + R2 e k d (τ s1E ( t )τ S 2 ) ) (62), a = a2 (1 − 0.5[C10 ] (63) (1 + R1 + R2 )

Energy Balance (flat radial profile of temperature) 1.5 0.3 ζ (T − TS ) aη1 (k11CC2 CCIB + k3CC 2 C6 + k 6 aη1CC 2CC 8 )(−∆H1 ) + + ak9η1CC 2 C H 2 )(− ∆H 9 ) − Uat (T − Tc) dT = SL−S SL − r dV E [ A + K IC 4 C IC 4 + K C 6−C10CC 6−C10 ]E

(64)

+ ((ak8CC1.25CC 6−C 8 / 1 + 4C H 2,l )(−∆H 8 ) / E ) w / wo

1.5 0.3 ζ SL−S (TSL − TS ) aη1 (k11CC2 CCIB + k3CC 2C6 + k6 aη1CC 2CC 8 )(−∆H1 ) + +ak9η1CC 2CH 2 )(−∆H 9 ) − Uat (T − Tc) dT = − dV D G [ A + K IC 4C IC 4 + K C 6−10CC 6−C10 ].G

(65)

+ ((ak8CC1.52CC 6−C 8 / 1 + 4CH 2,l )(−∆H 8 ) / G)w / wo

E =[

ρ l Cp s voR2 ρ l Cp s vo(1 + R1 + R2 ) ( ρ L Cpl + ] (67) ( ρ L Cpl + + n g Cp g )] (66); G = [ hl (1 / Fc − 1) hl (1 / Fc − 1) R2 = U SW ( D 2 − (1 .2 Dd2 ) / FSL (68) [T f [

ρ l Cp S (1 + R1 ) (

T2 =

1 − 1) Fc

ρ l Cp S

+ R1 ( ρ l Cpl + n g Cp g )] + R2 [ ρ l Cpl + (

ρ l Cp S

(1 + R1 + R2 )[ ρ l Cpl + (

1 − 1) Fc

1 − 1) Fc

]T4

(69)

+ ngCp g ]

where Fc is the solid content in the slurry; T2 = f (To ) and T4 = f (T f ) are functions, described by equations (67) and (68), which can be easily determined by a heat balance when the olefins conversion is fixed by designing targets. a0, a1 and a2 are the activity of catalyst fed to the reactor, activity at the inlet and outlet of the riser. Vr and VD are the riser and down-comer volumes that depend on gas hold up. They, after integration, become V1 and V2. Boundary conditions can be summarized by the following equations: Gas phase: V r = 0, PC 2= : PC22= ; PC 3= : PC23 ; PH 2 : PH2 2 ; T : T2 ; V r : V1 , PC 2= : PC42= ; PC 3= : PC43= , PH 2 : PH2 2T = T4

(70) Liquid:


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V r = 0, CC 2= : CC2 2= ; CC 3= : CC2 3 ; CH 2 : CH2 2 ; T : T2 ; V r : V1 , CC 2= : CC4 2= ; CC 3= : CC4 3= , CH 2 : CH2 2 ; T = T4 V D = 0, CC 2= : CC3 2= ; CC 3= : CC3 3 ; CH 2 : CH3 2 ; T : T2 ; V r : V1 , CC 2= : CC4 2= ; CC 3= : CC4 3= , CH 2 : CH4 2 ; T = T4

(71)

Solid: V r = 0, a = a2 T = T2 ; V r = V1 , a = a3 T = T3 (72), V D = 0, a = a3 , T = T3 ; V D = V2 , a = a4 , T = T4 (72)

the streams properties at the inlet are: Fi,0 , ao ,T0 ,σ 0 , ρi,0 µi,0 ) (fresh feeds plus recycled ones; they are well mixed with the downcomer (assumed) stream ( Fi , 4 , a4 , T4 ,σ 4 , ρ i , 4 µi , 4 ) ; the new riser inlet properties are calculated using the VLE correlations and molar contributions at the same operating pressure ( Fi , 2 , a2 , T2 , σ 2 , ρ i , 2 µ i , 2 ) . The PFRR model uses these properties to start the forward numerical calculation of the conversion. 2.2.3

Hydrodynamic of the STR. In the riser, slurry phase is pneumatically moved up by

the gas phase. Bubbles provide the necessary turbulence to obtain a high contact and high interfacial area between the phases and to enhance mass and heat transfer from the solid to the heat exchanger. In downcomer the slurry moves downward, and the reactions occurred in presence of dissolved hydrogen. The simulation model takes in account the effects of solid average particles diameter, density of slurry, viscosity and local solid gas and liquid concentrations. Accurate predictions of gas hold-up and slurry concentration are essential since they affect olefins conversions, selectivity and coke accumulation. Hydrodynamic parameters are summarized in Table 4. Empirical equations were previously obtained based on concepts developed by Suganuma and Yamanishi (cited in reference19); they were used to estimate gas hold up (76), bubble velocity (77), solid distribution along the riser reactor (78), average bubble size (79), effect of solid in slurry viscosity (80), and Cv coefficients (Eq. 81). Notice that values of slurry density, interphase surface tension and gas density affect the average the size of the bubbles at elevated pressures (1-2 MPa) and moderate temperatures (340-380 K.) The phase behavior within STR and the mixing effect are the result of a complex iteration of the four distinct areas used in the model (Inlet, Riser, Freeboard and Downcomer zones). Slurry apparent axial dispersion is calculated using equation (82) previously developed by Galiasso and Salva16 based in the work of Hikita and Kikukawa (cited in reference

19

.) Recycle ratio (R2) is calculated using slurry dispersion and degree of segregation using Eq. 8316. Carberry24 established the relation between slurry dispersion, number of CSTRs in series and the recycle ratio for their PFRR model, without reaction in the recycle; the model was extended by us to take in account conversions and heat of reaction in the downcomer (Eq. (84)). N-CSTR model is used here only to provide a physical reference of


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PFRR backmixing. In practice, recycle ratio (R1) is mainly dependent on gas superficial velocity, and slurry properties. Slurry density (ρsL) is calculated by mass contribution of the solid and liquid density; liquid density takes in account hydrocarbon vaporization and hydrogen solubility at equilibrium. Gas and liquid flow rates at the inlet (Fg, Fl mol/s) is used to determine superficial gas and liquid velocities (Ug, Ul). Plug-flow recycle reactor (PFRR) model, used to simulate the STR, was developed and validated with reactant conversion and degree of recycling for butane alkylation12. Numerical integration (RKF) of this model is relatively simple because recycle ratios depends only on gas flow rate parameter, degree of segregation and properties of the fluids that were determined in cold models. Interphase between bubbles and slurry is affected by the relative velocity of the phases (Eq. 85.) For spherical bubbles, shear forces are caused the variation in the slip velocity but, at sufficiently low Reynolds numbers, they are approximated by the standard Schiller-Naumann equation (cited in ref.19.) Our previous simulation work demonstrated that gas-liquid and liquid solid mass and heat transfer rates effects in reaction rate are negligible neverless they are included in the model; mass transfer at the interphase liquid-solid depend on slip velocity, which is a function of transport properties through the drag coefficient (Eq. 86.) Solid transported into the freeboard by bubble burst at the interphase are calculated by Eq. (87) and the freeboard solid distribution by Eq. (88), both determined in the cold model study17. Contribution of reactions in the freeboard are neglected. 2.2.4 Gas-liquid equilibrium in the STR. The set of mass and energy balance equations mentioned above depends on concentrations and temperature of the slurry and the solid. The vapor liquid concentration at the interface is obtained from thermodynamic equilibrium calculations (VLE). Isothermal flash calculation was performed to consider each component in the reactants and product mixture within the reactors. It provides equilibrium composition, equilibrium distribution ratio ( Ki (T,P,xi , yi ) , and partial vaporization. The simulation program includes correlation developed using a Chao-Seader25 thermodynamic model included in Peng Robinson state equation to evaluate the vapor phase; Hildebrand activity coefficient was used for liquid phase. The activity coefficient is further corrected assuming Pitzer-Curl corresponding state principle. Gas-liquid equilibrium are calculated using equation depicted in annex I (Eqs. 93 to 99.) The VLE empirical correlations were included in the simulation program (Fig. 8); they were developed in PRO/II®.


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2.3 Program for PFRR calculation The PFRR model simulation need the input of mass flowrates and properties of feeds, ratio of catalyst and recycles streams to the reactor, the expected level of conversions, the kinetic and thermodynamic parameters, the operating conditions for all equipment and the expected performance (conversion, temperature, degree of separation, and so on). To start with the simulation, global mass and heat balance for the whole process were performed to define the mass flowrates and temperature of the streams. Recycle flowrates are assumed and simulation calculate the rest of the input values need to solve the PFRR model and calculate the volume and dimensions. The program also calculates other equips sizes and costs (heat exchanger, recycle compressor, gas-liquid separation and pumps using conventional short-cut methods. The program simulated the effects of operation parameters in an iterative procedure, following the algorithm shown in Fig. 8. Ethylene conversion is used as convergence criteria for calculating the reactor volume; in other mode of operation the program determined the conversions, temperature and yields, at constant reactor volume. Thus, the program predicts the volumes of reaction zones (riser: V1; downcomer: V2) and freeboard volume (V3). When the program converged to the specified olefins conversions, it gives as output the following characteristic design and operation parameters: STR volume, temperatures, catalyst activity and concentrations along the reactor, recycle ratios (R) and the fluid dynamic parameter; the program also define the composition, mass flow rate and temperature of the gas phase and the slurry leaving the reactor. Finally, the program calculates the investment using the main equips dimensions or characteristic parameter and materials, as well as utilities consumption and the operating cost. The following are the hydrodynamic assumptions: (1)

Constant average bubble size along the axis of the reactor is calculated based on statistical coalescence and re-dispersion. That implied a constant number of bubbles. Nevertheless, the simulation model can use a variable size of bubbles point by point.

(2)

Ideal heat transfer at the bundles of tubes (flat radial profile of temperature).

(3)

Bubble are at the same temperature of the slurry phase along the reactor.


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Energy & Fuels

(4)

Reactions take place at the solid in the slurry. There is negligible amount of catalyst content in gas phase (cloud and wake).

(5)

Pressure drop is not included in the balances. There is less than 10% of pressure losses, a fact verified in the cold model study;

(6)

The system operates as close-open system; mixing of the slurry phase is represented by an PFRR with internal heat transfer model;

(7)

There is an axial profile of solid in the riser that depend on average particle size. The profile is taking in account by the w/wo correlation without solid segregation.

(8)

Slurry in downcomer zone of the reactor does not carry bubbles and move as plugflow. There is a flat axial profile of solid

(9)

The cone area well mixed slurry bubbling zone with negligible contribution to the conversion.

Table 5 depicts catalyst properties needed in the simulation and Table 6 the input parameters; Simulation model outputs are shown in Table 7 and product qualities in Table 8, as a reference. Base Case (BC) operating conditions are: 373 K, 1.2 MPa, H2/IB: 0.3 and H2/O: 2 molars; ethylene conversion 0.51; PH2: 0.9 MPa, dp: 50 microns Cat/HC: 8 % by wt. Fig. 11 shows the profile of reactant and products, coke and catalyst activity as a function of reactor volume for BC. Table 9 shows the effects of departing from base case operating conditions. Table 10 and 11 describe the effect of mass transfer and Fig. 12 the effect of inlet temperature. Table 10 and 11 present the effect of mass transfer and Fig. 12 shows the effect of backmixing and temperature. 2.4 Hydrocyclone-settler. The gas-liquid-solid separation was accomplish using two hydrocyclone-settler operating in series; they are designed to treat a combined gas-liquid-solid stream leaving the STR (see Fig. 8), with hot hydrogen. This special unit contact the stream flowing in tangentially in a cylindrical part and hot hydrogen bubbling from the cone (bottom) to achieve deep cyclonic separation of liquid and solids from gas and liquid. Notice that delta of pressure in this equip affect the internal recycle ratio (R) in STR. Design parameter used for HCS affect separation efficiencies; they are described in detail in reference26 and were calculated based on modified Savarosky27 concepts. ACS Paragon Plus Environment

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2.5 Economical evaluation Fig. 1, left and right side, shows an example of flow scheme (reaction zones, product separation and recycles) which was simulated to determine capital and operating costs. The scheme contains several loops of recycled stream, as well as gases and solid makeups and purges. The base case (BC) operating conditions were established with 100 % gas-liquid-solid separations in the HCS and other fractionation stages. The BC simulation study of the integrated alkylation- regeneration system used calculated efficiency for HCS and other separation stages, but the difference in cost respect to the case with complete separations are minor28. Original simulation of the entire process was performed by Rodriguez Dominguez (2004)28 that was updated by the current authors in 2014. Description of the economical evaluation of the whole process is provided in complementary information. Briefly, the evaluation of alkylate cost was performed by calculating the size of main equips using two ad hoc reactor programs (one for STR and another for FBR) to design the reaction zone, plus the use of PRO/II® (2008)29. The latter is used to calculate dimension of separations, cooling and heating, compression and recycling of gases and fractionation stages (short cut methods.) Heat and material balance of the process scheme were used to determine the operating cost of the reactors (Gulf Coast USA 2014) following the procedure described in, for example, Tourton et al. (2008)30, among others. Details of criteria used for equips design and economic evaluation of the integrated process scheme are depicted in complementary information. The integrated simulation of the process at different STR operating variables is iterative because there are several recycles streams; the only way to shows the complex effects of variables in the process is through a comparative alkylate costs evaluation respect to a Base Case (see Table 9 in second part of the paper13) at constant capital cost (sizes of equips.) The change in operating conditions affects the product yields, quality of alkylate and utilities consumption (catalyst, cooling water, hydrogen, and electricity.) Therefore, the cost of alkylate ($/(octane x m3)) at constant capital expenditure was determined for BC and for different operating conditions (Table 9.) The level of capital investment determine here was verified using main equips vendors quotations and then compared to an installed HF plant of similar size in 2014 (Gulf Coast.) To describe the performance of the slurry transport reactor-hydrocyclone (STR-HC) system,


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the discussion of results that follow is focused on explaining the catalyst characteristic and the data obtained in batch micro-reactor and pilot plant tests; then, the simulation of a commercial size reactor and the effects of operating variables are discussed. 3. RESULTS AND DISCUSSIONS 3.1 Catalysts. Table 1 shows catalyst physical properties of spent and regenerated catalysts. Fresh catalyst (D) particle size distribution ranged between 40 and 60 microns (95%). The solid contain 0.15 % by wt. of Pt, well dispersed on surface of [TiZrSSiOx], where there are several types of acidic clusters. Bronsted-Lewis sites present on surface of micro-mesoporous solids are mainly responsible of the alkylation activity and hydrogen activation takes place on Pt. The fresh catalyst has a superficial molar concentration of 5.4 % of Ti, 7.4% of Zr and 3.4% of S, a total acidity of 0.45 micromoles of pyridine/g and a Brönsted-to-Lewis acid site ratio of 0.16 (adsorbed pyridine at 383 K measured by IR.) Per our model, the alkylation reactions occurred in Lewis acid sites10,17. Fresh catalyst D adsorbed at equilibrium 2.2x10-2 of iC4, 1.5x 10-3 of dimethyl-butane (DMB) and 2.1x10-3 micromoles/g of Hexylcyclohexane (HCH), at 363 K, 1.4 MPa of hydrogen, on a surface area of 320 m2/g (determined by nitrogen adsorption). Catalyst has a pore volume of 0.41cm3/g, 10% of which is from micropores of 0.8 nm (average pore diameter) and 90% from mesopores with of 8.0 nm (average pore diameter obtained with N2 and Ar adsorption-desorption.) The amorphous tridimensional pore structure was developed by using template and by treating the assynthesized meso-Silicalite-type precursor with steam-ammonia; fresh catalyst has a tortuosity factor of 1.2, measured by benzene diffusion at 373 K. The blend of fresh plus regenerated catalyst D33 has an initial activity (a) of 0.8-085 (conversion of olefins into light alkylate at t=0, 353 K, 1.4 MPa.) Spent catalyst, which was withdraw from a CSTR pilot plant at the base case operating conditions (50 % olefins conversion), has lost 53% of surface area, 28% of pore volume (72% of micro-pores volume), 62% of acid sites, 67% of iC4 adsorption and 77% of olefins alkylation activity respect to fresh catalyst (see in Table 1, properties of D33 with 0.8% of coke). The coke on catalyst has an average hydrogen-to-carbon molar ratio (H/C) higher than one. It is homogenously distributed in the catalyst (TEM analyses) and contains 0.64% of CS2-soluble (373 K) coke, most of them in mesopores, and 0.18% of insoluble-coke in


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micropores. The catalyst tortuosity factor in D33 sample is 0.86, calculated from benzene diffusion. Hydrogen-regenerated catalyst, at the base case operation conditions

13,31

contains an average

0.15 % wt. of carbon, 97% as insoluble coke still located in micropores. It shows a reduction of 43 % in surface area, 8% in pore volume, 36% in acid sites, 12% in iC4 adsorption and 49% in activity respect to fresh catalyst. The tortuosity factor is 0.98. Regenerated catalyst was purged, and make-up of fresh catalyst added; resulting properties are depicted in Table 1. 3.2 Alkylation: Kinetics and deactivation. Batch reactor tests (Fig. 2a). Previous alkylation tests17 in a batch reactor demonstrated that TiZrSSiOx catalyst, in absence of Pt and hydrogen (T: 353 K, P: 1.4 MPa, N2), built up fast coke, and after 0.1 h on stream 1-2% of coke was accumulated on surface. When the catalyst contains Pt, and the reaction is operated in presence of hydrogen (batch reactor, T: 353 K a, P: 1.4 MPa, H2), the initial rate of alkylation is also high, but the catalyst deactivation is 10 times slower than in previous case. A fast hydrogenation and desorption of lights intermediates products occurred due to large concentration of hydrides on surface (see the value of initial activity at t=0, reported in Table 1 for spent and fresh plus regenerated catalyst used in the pilot plant.) Fig 3a (points) shows, as an example, that the rates of olefins conversion and coke built up decrease with alkylation contact times (Runs 1 → 4) and increase with temperature in the batch reactor. Each point is the average of three independent measures of conversions and the bracket represent the maximum deviation. Detail analyses of light alkylate demonstrated that it is composed by 90% of C6-C10 iso-paraffins, with a tri-isomers/di-isomer ratio of around 0.23, at ethylene and propylene conversions in the range of 40 to 60%. The simulation of the batch reactor uses eleven (lumped) kinetic reactions rates model (Table 2, equations 9-20). Mass balances equations used (depict in Annex I, Eqs. 98-110) were solved using a RungeKutta numerical method and the parameter adjusted by a Genetic algorithm. The result of the simulation, with adjusted kinetic rate parameters, are shown in Fig 3a (lines.) The simulation fit experimental results with deviation of ± 0.84% respect to the expected value and confirms the effects of temperatures, contact time and particle size used in the kinetic rates model.


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Carbon distribution is mainly dependent on the relative rate of primary alkylation reactions (C2= + [iC4+], C3= + [iC4+]), and re-alkylation (C2= + [iC6-7+], C3= + [iC6-7+]) because [iC6-iC7] intermediates are fast hydrogenated by the “activated” hydrides and desorb at rate q1k11; thus, primary alkylation intermediates suffer minor re-alkylation, dehydrogenation, dimerization and aromatization reactions17. However, longer the residence time, higher the number of adsorbed intermediaries that react with olefins or dimerize to produce high molecular weight products that do not desorb or desorb at very slow rates. Notice that higher the molecular weight of product, lower the rate of desorption (see values of q in Table 2, equation (20) for different carbon number.) This trend was experimentally being verify in batch studies, and it agrees with their energy of adsorption (see also DFT calculations of adsorption energies in reference17.) Dimers and naphthenic molecules ([C10-C16]++) do not desorb (q4k11=0) because they have multiple points of strong adsorptions, thus they form soluble-coke and then aged into insoluble-one. The effect of increasing particle sizes (shows in Fig. 3b, bars) is to reduce the reaction rates at constant operating conditions (Run 3, 353 K, 1.4 MPa.) The smaller the particle size, the higher is the amount of alkylate and the lower coke accumulation for the same contact time and total pressure. Similar tests were performed with samples containing different levels of insoluble and non- regenerable coke (from 0.15 to 0.3 % by wt.) Insoluble-coke is accumulated on catalyst after several passes (“ages”) through the alkylation-regeneration system. Fig. 3b dash-lines shows that selectivity to light alkylate decreases when particle size and coke increases. The comparison of predicted values with 88 experimental data points, obtained at different oprating conditions, particle size and level of coke, allows the estimation of catalyst deactivation (a: f(coke) and average catalyst efficiency factor (η: g(coke)) in the range of 40 to 60% of olefins conversions. These functions were then confirm using the pilot plant data (see below.) Fig. 3b (bars) also indicates that there is a small effect of adding dimethyl-butane (DMB and Dimethyl-Pentane (a weak adsorbed product of reaction) on the rate of olefins conversion and coke built-up. The same tests, but in presence of 0.1 hexyl-cyclohexane (HCH, one of the heavy intermediaries that built-up coke), shows a small reduction in alkylate production at the ACS Paragon Plus Environment

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same operating conditions. The rate of alkylate production is inversely proportional to the ratio HCH/iC4 (0.1, 0.2 or 0.3) used in the reactor at t=0. HCH was strongly adsorbed at t=0 and its presence reduces the number of sites available for alkylation reactions and increase the rate of coke production. Fig. 3b (bars) also shows small positive effect of hydrogen partial pressure on the rate of reaction in agreement with the proposed order for hydrogen. The batch reactor operates without bubbles-stripping effects, but solid and liquid are well mixed. Bubbles are very small (less than 1 mm) and well dispersed. Sampling at different level demonstrate that the slurry is well mixed reactor. Higher the temperature used lower the amounts of olefins and iC4, and higher the amount of hydrogen present in liquid phase. Gas hold up increase linearly with the temperature in the range 353-383 K). The effect of temperature in the kinetic rates of reactions takes in account the gas-liquid equilibrium The kinetic rates model (eqs. 1-8) allows us to establish the solid make-up and withdrawal needed by the pilot plant to operate at steady-state at different operating conditions (semicontinuous and continuous tests.) Pilot plant CSTR. Preliminary tests. Twenty-two tests without reaction were performed in the pilot plant reactor using regenerated catalyst and hexane to determine the effects of several fluid dynamic parameters in olefin conversions and selectivity. 0.8 L/D ratio is used for the slurryphase; gas is distributed at the bottom and flows up at linear velocity of 0.02-0.04 m/s. The measures of local delta of pressure and laser local bubble frequency have demonstrated the presence of tiny homogenously-distributed bubbles in the slurry phase. Bubbles of less than 0.03 m diameter was observed rising at 0.62 cm/s gas velocity (laser probe sensor), for this type of impeller and for this special gas-sparger design; global gas hold up (εg) is of 0.14 and global hydrocarbon vaporization of 45% at 353 K (1.4 MPa). Rate of stirring was set to 300 rpm; hold up measured by laser and solid hold up determined by sampling were similar at the bottom and near gas-liquid interphase level in the slurry phase, hence the reactor is well mixed due to double blade stirring effects. Pilot plant were operated in the range 4 to 20 % of solid content. The tests demonstrated that gas hold up measured in pilot plant is not dependent on particle size in the range of 50-170 microns. Gas hold up and hydrocarbon gas-liquid


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equilibrium changed with temperature and pressure, thus the slurry volume and the number of moles were calculated for each of operating condition to determine hydrocarbon concentrations. The correction introduces changes higher than 5% in conversion to hydrocarbon due to hydrocarbons vaporization and hydrogen solubility. Samples of (wet) silicalite containing reactant and product were analyzed by CS2 extraction to confirm the concentration in pores. The results demonstrate differences in concentration measured and calculated lower than 1.2 % at three temperatures. Bubbles carry-out gases (olefins, hydrogen and gaseous product of reaction) generating a stripping effect that help the rate of mass transfer (the same effect is observed in the riser of STR.) Seventeen preliminary reaction tests (1.4 MPa, 353 K, 0.2 h and 0.3 h of gas and solid residence time) in the pilot plant reactor have proved that there is no effect of gas feed rates in the range of 0.01-0.06m/s and that stirring rate (100-300 rpm) in olefins conversions and product distribution. Forty scouting tests performed at different Cat/HC and H2/HC ratios (P: 1.4 MPa) proved that olefins conversions are proportional to catalyst/oil ratio in the range of 4 to 12% and they are slightly dependent on hydrogen partial pressure (from 0.6 to 0.9 MPa). 50 Kg of D catalyst were prepared by the catalyst-vendors to perform the experiments. Continuous operation. Fig. 4 shows the olefins conversion into alkylate coke and light paraffins when it was operated in transient (due to change of operating variables) or in steady state. Different temperatures, residence times, catalyst particle sizes and solid throughputs were explored. This figure depicts, for simplicity, only one point of coke content per day of operation, but each point represents the average of six samples that were analyzed independently; the composition of products during transient operation allow us to verify that there was fast approach of steady state for different levels of conversions and temperatures. The analysis of product of reactions confirms that by increasing temperature, the undesired reactions that form coke and hydrogenate olefins are favored and light alkylate yield decreased. Any combination of temperature, residence time and catalyst throughput that convert olefins above 60%, increases more the yields of soluble coke and light paraffins than light alkylates. octane number were measured in 45 large samples of light alkylate, obtained in the range of 40 to 60% olefin conversion; octane number varied around 91-93 (Fig. 4) in these highly isomerized products.


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Fig. 5 depicts carbon distribution on alkylate at different levels of conversions in the range of 40 to 60% conversion obtained in the batch reactor. Left Y-axis shows the iC6 -iC9 content in light alkylate and right Y-axis the iC10+, C2+C3 and C4+C5 paraffins formed, as well as the weight of soluble coke on catalyst and the ratio of naphthenes / paraffins formed on light alkylate. The production of hydrocarbon C10 - C16 are relatively smaller than iC6 - iC9 (light alkylate) when the conversion is limited to 50%. Notice that C6 - C7 carbon number decrease and C8 - C9 increase when olefins conversions are increased due to the in-series re-alkylation and dimerization reactions. Higher the conversion higher the ratio the carbon number and the naphthenes / paraffins ratio (see equation [89] below.) Degree of isomerization (ratio tri + di / mono-isoparaffins) increases quasi linearly with the carbon number (Eq. 90 below.) The small variation in carbon number in a narrow range of conversion and the high degree of isomerization are indicative that the octane can be simulate by an empirical (simple) function of carbon number (Eq. 91.) The following equations were found to fit the analyses of 45 samples in the range of 40 to 60 % olefins conversions (carbon by weight): (Naphthenes) /Paraffins (-)) : 0.12 + 0.005C6 + 0.019C 7 + 0.026C8 + 0.033C9

(89)

(Degree of isomerization (-)) : 0.504 + 0.289C 6 + 0.454C 7 + 0.341C8 + 0.588C9

(90)

(Octane number (-)) : 33.6 + 0.89C6 + 0.93C7 + 0.95C8 + 0.93C9

(91)

Apparent kinetic rates expressions for this alkylation process obtained in the batch reactor were verified by the data of olefin conversion obtained in the CSTR. The program that simulate the steady state operation (adiabatic CSTR equations) confirm that primary alkylation on this catalyst follows an apparent first order or reaction for each of the hydrocarbons (olefins and intermediates.) Rate of heavy naphthenic hydrocarbons production depend on concentrations of adsorbed intermediary at order one, olefins at order 1.3 (lumped kinetic rate equations that represent six in series parallel equations), and it is inversely proportional to hydrogen partial pressure. Hydrogenation depend on concentration of olefins to order one and hydrogen partial pressure to order 0.5. Hence, the sets of apparent reaction rate constants (Table 2 and 3) demonstrate that they can represent well the results in the batch reactor and in CSTR when operates in the range of 40 to 60% conversion.


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The comparison of simulated vs. measured olefins conversion into light alkylate coke and paraffins, show in Fig. 6a, b and c, indicates that the CSTR model fits experimental data (pilot plant tests) with a deviation of ±0.65% in 95 % confidence range for all particle diameters. The model predicted average a and η with a deviation of ± 0.5 %, at 95% confidence range. Fig. 7a shows the comparison of delta of activity predicted and measured in steady state. Previous fundamental alkylation studies17 demonstrated that on this catalyst, containing Pt, there is higher rate of hydride transfer than in other zeolites used for alkylation7,9; that preserves the primary alkylation products and limit the soluble coke aging, but convert olefins into light paraffins. Hence there is a compromise in Pt content and hydrogen partial pressure. Primary alkylation reactions have lower activation energy than re-alkylation, hydrogenation and coke formation ones (Table 3). The mechanism of reactions proposed explain why there is abundant di- and tri-methyl-butanes (Table 8) and low naphthenes and aromatics content in the alkylate. There is very small amounts of desorbed cyclo-paraffinic hydrocarbons (C6 – C9) which are expected products of cyclization reactions. Less amounts of cycled compounds are formed on current catalyst than on other zeolitic catalysts9 and on ZrTiSOx catalyst17, due to its hydrogenation-isomerization capabilities and type of acid sites. Current apparent pseudo-homogenous pre-exponential factor (k’o in equations (9) to (18) Table 3), include the gas-liquid equilibrium constants, liquid-solid partition coefficients and the adsorption term ([ 1 + K 6,7 CC 6−C 7 + K 8,9,10CC 8−C10 ]) normally used in on-surface reaction heterogeneous model; the adsorption term is quasi-invariant in the range of 40-60 % olefins conversions used here. Therefore, by using homogenous kinetic expressions the simulation program fit well (±1.5 %, not shown) the empirical equations (89) to (91). Change in alkylate density is predicted based in composition. The analyses of the alkylate composition (for example those of Table 8), obtained in the pilot plant around 50% olefins conversion, shows small increase in average molecular weight and octane number and - reduction in liquid yields when temperature is increased, at similar level of olefins conversion and particle diameter. Equation (80) ws developed to predict octane number and alkylate cost (defined as $/octane m3) as a function of carbon number. Some C8


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isomers have the highest octane-number, therefore when average carbon number of alkylate increases, the octane number (calculated by contribution of species) slightly increases at expenses of volumetric yield of alkylate; hence the octane-m3 of the alkylate pass by maximum and then decreases due to the conversion of [iC6] and [iC7] (di- and tri-branches adsorbed intermediaries) into heavier di and tri-branched isomers (iC8-iC10). The spent catalyst leaving the pilot plant contains a distribution of soluble-coke content as a function of number of particle equivalent to one well mixed compartment (ECSTR function), as expected. The kinetic and deactivation rates model is used to design and simulate a STR for a commercial scale process. The model fit experimental data both reactor that represent the extreme case in fluid dynamic; thus, it is independent of the backmixing of the equip used. 3.3 STR design and operating variables. The simulation of a STR-HCS system used a PFRR model based on a fluid-dynamic model and the apparent kinetic and deactivation rates model mentioned above. The PFRR model have been previously developed in a cold model

16

(without reaction) and it consider the

reactor composed by four zones: 1) a gas-liquid-solid distribution (inlet) 2) a riser-vertical heat exchanger 3) a downcomer-vertical heat exchanger and 4) the freeboard. STR is designed for low gas linear velocities (0.18-0.3 m/s) in the riser and (0.06-0.08 m/s) in downcomer that produce a smooth solid-liquid internal re-circulation velocity. Bubbles of gas are not entrained into the downcomer at this range of gas linear velocities. The bundle of tubes (heat exchanger) separated the riser and the downcomer (see details in Fig. 8.) Reactor and auxiliaries equips are designed for a Base Case operating conditions described below. 3.3.1 STR-HCS. Base Case (BC) design and simulation. The STR operating conditions selected for the BC design are: •

373 K, 1.4 MPa, 0.3 H2/iC4, 2 H2/O: 2 (molar ratios), 0.51 ethylene conversion, 0.9 MPa PH2, 50-microns dp, 8% by wt. Cat/HC ratio.


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Energy & Fuels

•

HCS is assumed to ideally separate gas-liquid-solid, regeneration stage (Base Case) operating conditions are: 533 K, 1.4 MPa, 0.2 h of Tg , 0.05 h of TS , 200 Pa of delta pressure

•

The regeneration of catalyst is performed at 533 K, 1.4 MPa, Tg: 0.2 with solid stripped at 473 K by hot hydrogen. That generated a catalyst with 0.2 % by wt. of coke that contain 0.05 % by wt. of insoluble coke.

•

The fractionations stages (gas-liquid separator and columns) are designed to obtain 89% of ethylene, 87% of propylene and 93% of isobutane recovery that are recycled to the STR.

The global process produces 1000 cubic meter per day of alkylate. All the feeds throughputs are mention here in metric tons per hour (m t/h). At inlet, there are 12.91 of C2=, 21.3 of C3=, 113.8 of iC4, 34 of others inert hydrocarbons (n-C3, n-C4), 10.83 of the regenerated catalyst and 0.05 of fresh catalysts (make up), and 3.65 mt/h of hydrogen (0.8% molar purity). Length-to-diameter ratio of the slurry zone is set to 7, global liquid residence time to 0.31 h for 50% olefins conversion, gas velocity to 0.18 m/s and slurry velocity in the riser to 0.014 m/s. Riser diameter (Dc) is 1.93 m and reactor diameter (Dr) 2.2 m. There is gas distributor with 124 holes of 2.1 cm, two-gas solid cyclones and a two-pass bundle of 800 tubes (2” diameter) used as heat exchanger along the axis and cooled by water (300 K). The input for the simulation are resumed in Table 6. They are introduced into the simulation program (Fig.10) that solve the differential equations using a numerical method. Subroutines calculates the dimensions of the reactor, which are reported in Table 7, the dimensions of other equips (heat exchangers, compressor gas-liquid separator), utilities needed, capital and operating costs. The reactor program converged in 12 CPU for the BC operating conditions. Table 7 resume the main results of the simulation; 51% or C2 and 60% of C3 global olefins conversion were obtained by using a volume of 57 m3 in the slurry section (V1(gas)+V2(liquid), 14 m3 in disengaging section (V3) and 3.5 m3 at the cone (inlet). In downcomer, conversion of C2= and C3= olefins increase from 51% and 60% to 59 and 69 % respectively.


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L/D ratio of the two reaction zones were calculated to fulfil the specified gas linear velocities. The riser has 10% of gas hold up and 81% of liquid hold; average bubbles size is of 0.08m diameter; similar bubble size was reported by Luo et al.32 in different reaction system. The model predicts an internal slurry recycle ratio of 0.35 based on the values of gas, solids, and liquid flow rates in the riser and properties of the phases at the inlet of the reactor. The linear velocity in the riser was selected to produce a stable disengaging interphase16. Fig. 11 shows the profile of reactant catalyst activity and coke as a function of riser volume. Similar profiles were obtained in downcomer but with smaller slops in olefins conversion, catalyst deactivation and coke built up. Effect of fluid dynamic parameters. Gas holdup in the riser is slightly augmented with an increase in gas or liquid velocity but is reduced by increasing solids loading or slurry density and viscosity. Apparent liquid circulation rate increases as the gas or liquid velocity increases without modifying the global gas hold up, almost in the range of gas linear velocity of 0.2 0.3 m/s. Slop of profile of activity, coke and reactants depicted in Fig. 11 decrease by increasing gas velocity in the riser; the higher is the hydrogen flowrate, the higher the recycle ratio (backmixing) and the smaller liquid residence time in the riser and in downcomer. Hence, gas linear velocity change activity, selectivity and catalyst deactivation by changing the backmixing. Notice that there is an optimal solid-load and gas velocity in the slurry at which slurry does not carry-down bubbles in downcomer zone but still produce important global backmixing. Riser gas holdup decreases with pressure and increases with temperature as it is shown in our empirical correlations (77)16. Bubble size is 9% reduced by increasing solid content from 4 to 8% solid concentration and 12 % by increasing average particle size from 50 to 150 microns, at gas linear flow rate of 0.2-0.3 m/s and slurry velocity 3-5 times above settling velocity. Previous cold model studies have shown that solids are uniformly distributed in the riser and downcomer reactors17 when slurry viscosity is in the range of 0.24-0.25 Pa.s, density between 600-800 kg/m3 and surface tension around 0.02-0.22 N/m, values used in the STR design. Slurry phase, a BC operating conditions, is lighter, less viscous and with higher surface tension than those commercially tested by Sasol in F-T synthesis reactor (T: 513K, P: 3 MPa, uog: 0.2 m/s, 808 k/m3, 0.45 MPa.s of viscosity and 0.018 N/m of surface tension for liquid.) ACS Paragon Plus Environment

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Bubble size is expected to grow slowly along the axis and to develop an average diameter below the maximum size that breakup32. Freeboard. Bubbles burst at the interphase at the rate of 5.5 bubbles/s projecting gas, drop of liquid and solid present in clouds-wakes into the freeboard zone. There are 0.018 m t/h of solid particles and 0.18 m t/h of liquid drops (ddrops~1 mm) projected into the freeboard based on gas and liquid linear velocity used; most of liquid drops are soon settled back (89% at 0.4 m high) into the interphase producing a short cloud region of 0.08-0.12 m over the interphase (value based on laser diffraction measures in pilot plant.) Bubbles are formed at gas-sparger region and burst at the interphase of dense bed producing a total pressure oscillation of 0.22 KPa. In total 0.618 m t/h of gases are transported into the cyclones (with a molar composition of: 0.527 H2, 0.203 C1, 0.152 C2 and 0.077I C4) at gas linear velocity of 0.04 m/s; this velocity limits up to 2% the amounts of coarse solids (40-50 microns) transported into the cyclones. At the outlet of the reactor, the solid content in gases is only 0.0001 m t/h, 92% of them are fines (1-10 microns diameter of particles.) The rest of solids come down through the cyclone legs into the slurry phase. 47.4 m t/h of slurry are internally recycled back though downcomer from the interphase to the gas-sparger zone. Gases at inlet. Gases and fines of solid leaving the top of the reactor pass through a filter, they are cooled, and condensed hydrocarbons are separated and send to the fractionation stage. 0.11 % of total gases are purged and equal amount of fresh makeup hydrogen added (see process scheme in Fig. 1); 0.012 m t/h of gas is purged, and the 0.016 mt/h of fresh hydrogen added at the inlet of STR, to keep constant the hydrogen partial pressure in the recycle stream at this level of conversion. Purged gas stream is sent to PSA unit where hydrogen and HC are recovered. 136.4 m t/h of slurry are sent to the HCS system to recover gas, solid and liquid.; despite of bundles of tubes heat transfer in the reactor, downcomer temperature increases from 393K to 408 K in BC simulation. Catalyst deactivation and solid makeup. Products of STR. Deactivation, due to coke, increases along the riser and downcomer; the rate of deactivation increases with temperature and decreases with the hydrogen partial pressure (reactions (5) and (6)). At BC operating conditions there is a moderate accumulation of coke due to low inlet temperature and small delta of temperatures along reactor (delta outlet-inlet is lower than 12 K). Pilot plant at similar ACS Paragon Plus Environment

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operating conditions shows that 77% of the CS2-soluble hydrocarbon are naphthenic type (65% of then are alkyl cyclohexanes and 12% alkyl-decalines) and the other 23% are isoparaffines (Table 1). Naphthenic coke compounds content on soluble coke increases only 8% with conversion in the range of 50 to 60% used here. The 96% of CS2-soluble compounds are “regenerable” by hydrogen13 and the 4% aged into insoluble coke. Therefore, 0.46 m t/day at of no-regenerable coke at the base case operating conditions is accumulated (between 0.02 to 0.04 % by wt. per pass) by passing throughout alkylation-regeneration stages. Thus, to maintain the activity level in the reactor, and to compensate for “fines” formation (0.001 m t/h), 0.048 m t/h of regenerated catalyst is removed from the system and 0.050 m t/d of fresh catalyst is added. Higher the purge and make up, higher the average activity at the inlet of STR and higher the operating costs. Despite diffusional rates control of reactions, TEM analyses (not shown) of spent catalyst suggest a quasi-homogeneous distribution of carbon along the radius of the particle, in agreement with the mechanism of “regenerable” coke builtup via adsorbed species at the intersection of meso-micropores structure

17,18

. Due to the

excess of iC4 and the high mass transfer rate in mesopores, the activated-adsorbed intermediaries of reactions are well-disperse on surface. Coke concentration on catalyst surface (Fig. 11, right Y-axis), increases with residence time and recycle ratio, as well as catalyst deactivation; nevertheless, the rate of deactivation decreases along the reactor due to lower and lower olefins concentration (in-series reactions) in riser and downcomer zones. The solid along the STR accumulates 0.053 m t/h of carbon on catalyst to attain to 0.58% on catalyst at the outlet (Table 7, BC operating conditions.) The mixture of hydrogen-regenerated plus fresh catalysts, added at the reactor inlet, has moderate alkylation activity (a ~ 0.8, Table 1). At the riser-inlet, coke content 0.2-0.25 wt. % is the result of blending inlet catalyst with 0.13 % and donwcomer-solids with 0.8% of coke; this spent catalyst internally recycled (with a: 0.23) also produces a reduction of inlet average catalyst activity going into the riser. Base case cost evaluation. Dimensions of all equips (STR-HCS-FBR-Fractionation) were established for the base case as well as their operating costs (see complementary information.) The capital investment of the process is around 59 Millon dollar (2014), 33% of them if for the alkylation stage. The global cost of utilities (catalyst, steam, water, electricity) is 7.4 $/h,


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45 % of them in the alkylation stages. The alkylate cost for the base case is 25.01 $/(m3octane) estimated for a plant located at Gulf Coast USA, 2014 (see Table 9 in second part of this paper13). 48% of alkylate cost is associated to the alkylation plant. The 54% of contribution in this cost is due, other than feeds costs, to catalyst consumption, and only 7% is due to capital recovery expenses. Alkylate cost determined in this study for BC is 30% lower, than the alkylate market-value in 2014 used as a reference. Nevertheless, here only the delta of cost respect to BC is used. Notice that alkylate (Table 8), mainly compose by C6 - C9 isoparaffins, is lighter than HF-alkylate, and has only 93.6 octane number, factors that reduce its contribution to the gasoline pool value. Net back calculation of light alkylates estimated price (2014, Gulf Coast, USA, gasoline pool of a high conversion refinery), indicate a 10% lower octane-m3 value than conventional HF-alkylate. Beside the base case simulation, the effects of several operational parameters were also simulated (residence time, inlet temperature, reactant and catalyst ratio and catalyst particle diameter) with the same program but at constant size of equips and feeds throughput The cost of alkylate production was evaluated and the deviation from base case analyzed in term of alkylate process. 3.3.3 Effect of gas and slurry residence time in the STR. Profiles of olefins, isobutane, hydrogen (in kg mol/m3) and activity (-) of the catalyst in the riser and downcomer are shown in Fig. 11, left Y-axis, and coke formation in Fig. 11, right Y-axis for the BC. These profiles depend on gas and liquid flow rates, among other parameters. For instance, 15% increase in gas inlet velocity produce 6% increase in recycle ratio (R) and 2% reduction in olefins conversions (at constant reactor volume, liquid throughput and similar regime of bubbles in the riser and downcomer.) The increase in gas velocity inside the cylindrical bundle of tubes (riser) only augment 3% of gas hold up and produces 1.5% change in wi/wo, along the axis of the riser for 50-microns catalyst particles. The change in gas velocity slightly increases light alkylate yields, with small reduction in coke and olefins hydrogenation; this increase in gas velocity results in higher operating cost in the STR due to the need of additional recycle of gases, energy consumption and catalyst attrition. This change in gas flowrate does not affect the HCS stage.


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Table 9, second third and fourth rows summarize the effects of using different slurry residence time than BC, at constant reactor volume and gas throughput. The second of rows depicts the BC production of alkylate-1 (reactions 1, 3 and 5), alkylate-2 (reactions 2, 4 and 6), coke (reaction 7) and paraffins (reactions 8 and 9) and cost of alkylate.; for example, the third of rows depicts the effect of increasing the feeds throughput by 20% at constant STR volume. This change reduces 9.25% the yield of alkylate-1, 7.62 % of alkylate-2, 40% of coke and 9% of paraffins production respect to base case. In addition, it increases 4 % STR recirculation ratio (R). All of this effect results in 3.2% higher operating cost than BC, due to the need of additional product heating, cooling and separation; that happen despite that the change reduces 2.5% the catalyst regeneration cost13; globally the change augment 14.6% the cost of alkylate ($/(m3 octane)) respect to the base case. The fourth of rows shows that by rising the hydrocarbon residence time by 20% the cost of alkylate also increases 10.1%. Therefore, the global slurry residence time used for BC case is close to the optimum for these sizes of equips. 3.3.4 Effect of feeds flowrates. Activity, selectivity to alkylate and catalyst deactivation are sensible to the profile of temperature in the STR. Heat removal, gas and solid inlet temperature and internal recycle ratio are the parameter that control the temperature along the reactor for this highly exothermic reaction. The best selectivity always occurred for the flatter profile of temperature along the reactor (CSTR behavior). At base case operating conditions, the internal bundle of tubes (heat exchanger) remove 30% of the heat generated by reactions; the rest of heat of reaction is used to develop the profile of temperature. 3.3.5 Effect of inlet temperature. Table 9, fourth of rows, shows the effects of rising 8 K the inlet temperature respect to the base case (at constant heat removal by the heat exchanger.) The increase of inlet temperature generated 8.6% higher yields in alkylate-1 (from ethylene) and 13.9% in alkylate-2 (from propylene), 35.4% in light paraffin and 184.1% in coke; thus, this change of inlet temperature increases 40% the catalyst deactivation. The relative values of changes agree with of activation energies (Table 3) of the reactions involved. In addition, this change affects the fluid dynamics because reduces 1.8% de gas holdup, 2.4 % the recycle ratio and 1.5% the average bubble size, and increases 5 K the outlet temperature and 2.4% the gas-solid-liquid efficiency in separation stage (HCS)26. Left Y-axis of Fig. 12 shows an


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example of the effects of increasing 10 K the inlet temperature in the profile of activity along the riser and downcomer; Fig 12, right Y-axis depict delta of temperature in the riser and downcomer as a function of volumes; this change in inlet temperature augments 2 K the differential of temperature inlet-outlet, increases 5% olefins conversion and 21 % the cost of alkylate. The fifth of rows in the Table 9 shows that reduction of 10 the inlet temperature decreases (-4.8%) the cost of alkylate. The simulation results indicated the benefits of selecting as lower as possible inlet temperature than the BC, despite the larger reactor volume required for the same conversion. The results of the simulation, at different inlet temperature at the riser (343, 353 and 363 K) and different reactor volume (constant olefins conversion of 50%), indicate that at 343 K there is the best selectivity toward alkylate with the lowest catalyst deactivation and alkylate cost, despite of the 3% higher reactor volume. Notice that temperature at the inlet of riser (see detail in Fig. 8) is the result of mixing a hightemperature catalyst recycled from regeneration stage (~450 K), with isobutene, olefins and hydrogen streams (293 K) and with the slurry internally recycled (R: 0.3, T: 463 K). The inlet temperature is controlled by catalyst stream. The simulation confirm that the additional capital cost need is less important than the optimal use of olefins (selectivity.) 3.3.6 Effect of radial profile of temperature. STR cold model studies16 demonstrated that inlet conical section and the (sparger zone) operates as well mixing compartment for gas, solid and liquid; it plays an important role in homogenizing the inlet temperature at the bottom of the riser. Above this point and because of heat transfer, a radial profile of temperature starts to be developed in the riser despite the intense radial backmixing produces by the rising bubbles. In the simulation model, the mentioned heat balance for the riser and downcomer zone (eqs. (62) and (63), with constant value for U in Table 6), have been modified to consider a radial distribution of temperature near the bundle of tubes. Thus, the improved reactor model included a parabolic radial profile of temperature near the heat exchanger tubes, induced by a small change in linear velocity there. Equations (92-93) were used instead of (62-63) to calculate heat transfer. Hence, delta of temperature between the riser and the cooling bundle (TSL-Tc) of tubes were calculated along the reactor using a


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parabolic function of radial position. The equations were solved using an adaptive step-size control along the radius of the riser and dowcomer. ζ (T − TS ) dT = SL − S SL − dV r E

a η 1 ( k11C C 2 C C IB + k 3 C C1.52 C 6 + k 6 a η 1C C 2 C C 8 )( − ∆ H 1 ) + + ak 9η 1C C 2 C H0.32 )( − ∆ H 9 ) − 0 . 02 Pe (

d r 1 .6 2 r dT ) (1 − ( ) 2 dc dr dz

[ A + K IC 4 C IC 4 + K C 6 − C 10 C C 6 − C 10 ]E

+ (( ak 8 C C1.25 C C 6 − C 8 / 1 + 4C H 2 , l )( − ∆ H 8 ) / E ) w / wo

(92) ζ (T − TS ) dT = SL − S SL − dV D G

a η 1 ( k11C C 2 C C IB + k 3 C C1.52 C 6 + k 6 a η 1C C 2 C C 8 )( − ∆ H 1 ) + + ak 9η 1C C 2 C H0.32 )( − ∆ H 9 ) − 0 .02 Pe (

d r 1 .6 2 r dT ) (1 − ( ) 2 dc dr dz

(93)

[ A + K IC 4 C IC 4 + K C 6 −10 C C 6 − C 10 ].G

+ (( ak 8 C C1.25 C C 6 − C 8 / 1 + 4 C H 2 , l )( − ∆ H 8 ) / G ) w / wo

The improvement in heat transfer model reduces in half the speed of the simulation and produced less than 2-3 % changes in STR olefins conversions respect to BC. The preliminary simulation found that, if the delta of temperature (TSL-Tc) is lower than 15 K and the heat to be transfer limited to 30% of total heat generated by the reaction, olefins conversion predicted with radial profile only decrease 2.5% olefins conversion respect to the model that use flat radial temperature. Since temperature is so important in activity selectivity and deactivation, further experimental work should be concentrated to develop an improved heat transfer model 3.3.7 Effect of particle size and mass transfer limitations, STR-HC. External mass transfer. The values reported in Table 6 for mass transfer constants KGL and KL-S were used to calculate the effects of gas-liquid and liquid-solid mass transfer of olefins in

the reaction rates and ultimately in the cost of alkylate. These resistances where group as an external mass transfer in the simulation. Results of the simulation (Table 10), demonstrated that the rates of external mass transfer are not controlling the reactions rates for the three particle sizes studied here; their inclusion in the model produces less than 1% change in the cost of alkylate. Clearly, only mass transfer in pores controls alkylation reactions in the STR at this level of conversion, backmixing and temperature. Calculation of heat transfer rate (not shown) confirm that there is no heat transfer limitation of reactions at the liquid-solid interphase and in pores despite the high heat of reaction involved. Mass transfer in pores. Catalyst particle efficiency varied from 0.92 (35% conversion at the riser inlet) to 0.99 (0.5 % conversion riser top) for 50-microns particles diameter and BC


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operating conditions; catalyst particle efficiency varied quasi-linearly with coke content, increase with particle diameter and temperature. Internal mass transfer hindrance is one of the biggest setbacks of solid alkylation, leading to catalyst deactivation and low alkylate yields. Many author, for example Xiao and Wei23 (1992) discussed the diffusion (in gas phase) of heavy iso-paraffins in ZSM-5. They mention that the rate of diffusion decreases with the diameter of the molecules (degree of isomerization) but not with the length of the molecules. Still is unclear from previous published research and our own results18 how far desorption of products (soluble coke) in mesopores affect the rate of counter-diffusion (MS) in mesopores filled with liquid. Calculation shows that restricted diffusion in micropores only produces additional insolublecoke since most of micropores are filled with strong adsorbed reactant or blocked. Fig. 3b, dashed lines, show that the use of either 100 or 150 microns particle diameter instead of 50 microns reduce the rate of olefins conversion in a batch reactor. Larger the particle diameter, longer the path of diffusion for similar pore structure, and lower the reaction rates. The values of rate of reactions measured in batch reactor at t=0, with different amounts of coke on mesopores, allow the calculation of an effective diffusion coefficient in mesopores. Simulation obtain the effective particle diffusivity for DMB; for example, Deff,

DMB

was

0.29x10-9 cm2/s for carbon content of 0.14 wt % (D33 inlet of STR) and 0.24x10DMB-9 cm2/s for carbon content of 0.81 wt. % (D43, outlet of STR, Table 5). It was noticed before that the rate of desorption from different acid sites present on surface change with the strength of acid sites. The stronger the acid sites occupied for DMB or light alkylate, the lower is Deff . The number of strong acid sites decrease with soluble coke content. Figure 4 depicts the effects of particle size on olefins conversion in CSTR long-term tests operating at around 50% conversion. Notice that in CSTR the alkylation occurred in presence of large amount of iC4 and hydrogen on surface and small amounts of olefins; for example, at 50% of olefins conversion, there are between 3.0 to 5.8 moles of iC4 and 5 to 10 moles of hydrogen per mole of olefins in liquid phase; that generate a low concentration of olefins on surface, but high amounts of intermediates on surface. The effective diffusivity coefficient was also calculated for the CSTR data. Deff shows small difference as a function of coke


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content; in this narrow range of conversion, and the value is like those measured in batch reactor at the same level of coke. Make up of catalyst fed to the STR (D33, coke: 0.12 % by wt.) with 50 microns particle diameter has 60% of accessible micropores of 0.8 nm. There is 31 % area in mesopores of 9 nm available for adsorption. Surface area and pore volume (measured by N2) are 48% and 12% reduced in mesopores by accumulating 0.52 wt. % of soluble coke (D34.) This accumulation of coke reduced the filling of pores (vacuumed) by adsorbents: 19% for iC4, 8 % for DMB and 9% for HCH. IC4 can diffuse into micropores while DMB and HCH cannot18. Clearly soluble-coke is progressively deposited in mesopores (on mesopores walls and at the intersection meso-micropores) along the reactor, limiting the access and desorption from micropores and slightly reducing the particle efficiency factor. For example, at 50% conversion, this reduction is: 5 % for 50 microns, 7 % for 100 microns and 12 % 150 microns. Plotting η values (reaction 1 and 2) against particle diameter (not shown) in the range of 4060 % olefins conversion indicates that there is a quasi linear dependency of the particle efficiency factor with carbon content. Therefore, PFRR model uses a lineal correlation of η as a function of coke content to short the time needed for convergence. Table 1 shows that fresh plus regenerated catalyst has a tortuosity factor of 0.9, measured by benzene diffusion in gas phase, while spent catalyst has a tortuosity of 0.86 due to changes in the pores structure. Soluble-coke accumulation depends on how fast the alkylate intermediates are formed, hydrogenated (qik11) by hydrides and desorb. Hydrogen spillover help in some extend to reduce soluble coke built up, but it also converts olefins into paraffins, that affect the cost of alkylate. Notice that heavier the intermediate, slower the rate of hydrogenation (just compare k11 versus k7 and k8 (Table 3). Table 8, second and fifth columns, depicts the analyses of the

product at 375 K for 50 and 150 microns particle sizes at similar level of olefins conversion in samples of pilot plant. Diffusional control increases the cyclo-paraffins content and the octane number of the heavier alkylate produced. Soluble coke is formed during alkylation and then most of then it is regenerated by hydrogen. All of alkylation products formed inside micropores is or blocked by adsorption at the mouth or diffuse at extremely low rate (10-12 cm2/s). There, they are mainly converted into insoluble


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coke, which cannot be regenerated, and produce a waste of olefins and iC4. In practice after several cycles alkylation-regeneration only mesopores are useful to produce alkylate. Impact of particle diameter used. The change in particle size from 50 to 150 microns increases the reactor volume needed from 57 to 78 m3 at constant alkylate production (Table 11). Recycle ratio (R) and average bubble size are reduced in 4.5% and 2.8 % respectively, delta of temperature augmented in 5 K and coke content to 0.88 % by wt. on spent catalyst; that reduced activity in 40% and slightly increase octane number. The results of simulating the effect of particle size at constant reactor volume and other operating conditions are shown in Table 11. Observe that by increasing particle size from 50 to 150 microns produces 19% and 34% reduction in alkylate-1 and Alkylate-2, and 300% and increases coke and light paraffins yields in 10%. Notice that the PFRR program calculated point-by-point the solid content using a correlation that depends on particle diameter (equ.78.) Higher the particle size, higher is the solid content at the bottom of reactor and lower at the top (with quasi-linear profile of solid distribution.) However, the increase in particle size improve the separation efficiency in the HCS from 98 to 99.9%27 and reduce the volume of disengaging zone (V3) in the STR from 8 to 5 m3 due to lower carry-over at the same throughput of catalyst and gas flowrate. The increase in particle size also improve cyclones separation efficiency in the regeneration stage (see part II13.) The cost of producing 90-100 or 140-160 microns catalyst particle diameter increases 4 and 12 % respect to the cost of 40-60 microns. In summary, the use of 90-110 particle diameter instead of 40-60 microns in the process increases 25% the cost of alkylate ($/octane m3), despite of the former lower cost. Therefore, there is no interest in using higher particle size than 40-60 microns. Notice that catalyst cost represents the 54% of operating cost, not including the olefins cost. Particle size distribution (PaSD) also introduces important effects in fluid dynamics. When a broad particle size distribution is used, the smallest particle sizes are carried-over into the cyclones and HCS, rising the operating costs or the volume of reactor but not producing significant effects in fluid dynamic. PaSD also affect the regeneration stage (see part II13.) However, the use of narrow particle distribution increases the cost of producing the catalyst; for example: 30-70 microns cost 12% less than 40-60 microns.


Energy & Fuels 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Attrition of catalyst is another factor that affect the cost; fines of catalyst mainly affects volume of disengaging zone and catalyst make up (operating cost.) Fines particles increases entrainments and slightly the catalyst efficiency; to compensate the attrition additional fresh catalyst need to be added. PtZrTiSOx/SiO2 catalyst is based on hard-to-brake particles of siliceous oxide treated with steam-ammonia; consequently, only 0.0012%/h (0.027 m t/day) of fine particles was formed in pilot plant, which is a conservative value due to the impeller effects in breaking the solid; lower attrition is expected in a STR-HCS and FBR. Fines particles, collect at the bag-house, are sold to recover Pt in commercial size plant. Catalyst deactivation model. Catalyst deactivation model a: f(coke, Eq 63) was developed using the batch reactor tests (88 data points) at different operating conditions; the simulation of the CSTR predicts well (with deviation of ±1.3%) the amounts of coke formed in mesopores for these three particle sizes, two level of coke and three temperatures (see Fig. 7a). Therefore, the model is independent of the reactor used to measure the deactivation and can be applied to PFRR simulation. The PFRR simulation model uses the deactivation model (eq. 63) and the empirical equations (89-91) to predict the amount of alkylate, the carbon distribution and the octane number different level of operating conditions and degree of backmixing. 3.3.6 Effect of reactant ratios and catalyst concentration on conversion and selectivity of the STR. Hydrogen partial pressure. Hydrogen concentrations at the catalytic surface only change 2% with 10% increase in H2/HC ratio (achieved by increasing hydrogen purity at constant olefin and gas flowrate.) This change in H2/HC produces a negligible change in selectivity due the small order of reaction respect to hydrogen; for example, 10% increase in partial pressure respect to the BC produces minor ( gases flow rate > hydrogen partial pressure;

•

The lowest cost of alkylate in the STR-HCS is obtained at around 50 % of olefins conversion using 1.4 MPa of total pressure. The best sets of operating variables are: 333 K of inlet temperature, 0.2: O/iC4, 0:C2/C3, 0.8 Cat/HC and 0.03 H2/HC ratios, and for HCS the used of 1.1 mt/h of Hydrogen at 533 K.

•

PFRR simulation demonstrate that the minimum catalyst make-up needed is obtain for soluble-coke built up of 0.51 % in mesopores.

•

The value of the best cost of alkylate is mainly affected by the relative cost of olefins,


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catalyst and hydrogen and less by labor and capital recovery;

5. Acknowledgments The authors want to acknowledge the support of Simon Bolivar University, economical support of FONACIT of Venezuela and HyPro Co, the experimental work of Jose Andretti Salva and Juan R. Peretti, whom obtained the kinetic data, and the advice of Prof. P. Raminosof.

6. REFERENCES 1. de Jong K.P., Mesters C.M.A.M., Peferoen D.G.R., van Brugge P.T.M., de Groot C.; “Paraffin alkylation using zeolite catalysts in a slurry reactor: Chemical engineering principles to extend catalyst lifetime Chemical Reaction Engineering: From Fundamentals to Commercial Plants and Products Volume 51 (10), (1996) 2053–2060 2.

Corma A. “Inorganic solid acids and their use in acid-catalyzed hydrocarbon reactions”; Chem. Rev., 95 (3) (1995), 559–614

3. Feller A, Lercher J. “Chemistry and Technology of Isobutane/Alkene Alkylation Catalyzed by Liquid and Solid Acids” Advances in Catalysis 48, (2004) 229–295 4. Pater J., Cardona F., Canaff C., Gnep N.S., Szabo G., Guisnet M., “Alkylation of isobutane with 2-butene over a hfau zeolite composition of coke and deactivating effect”; Ind. Eng. Chen. Res. 38, (1999) 3822-3833. 5.

Weitkamp J., Maixner S., “Isobutane/butene alkylation on a LaNaY zeolite. Characterization of carbonaceous deposits by CP/MAS 13C NMR spectroscopy”; Zeolites 7, (1887) 6-8

6.

Flego C., Kiricsi I., Parker W.O. Jr., Clerici M.G., Spectroscopic studies of LaHY-FAU catalyst deactivation in the alkylation of isobutane with 1-butene”; Appl. Catal. A 124, (1995) 107.

7.

Feller A., Barth J.O., Guzman A., Zuazo I., Lercher J.A., “Deactivation Patways in zeolite –catalyst isobutene –butene alkylation”; J. of Catal. 220, (2003) 192-196.

8. Nivarthy, G. S., He, Y. J., Seshan, K., Lercher, J. A., “Elementary mechanistic steps and the influence of process variables in isobutane alkylation over HBEA”; Journal of Catalysis 176, (1998) 192-203


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9. Martinis J. M., Froment G.F., “Alkylation on Solid Acids. Part 1. Experimental Investigation of Catalyst Deactivation”; Ind. Eng. Chem. Res., 45, (2006) 940-953. 10. Galiasso Tailleur R. and Andretti Salva J, “1-butene alkylation reaction on PtSO4/TiZrOy catalyst” Preprint Aiche Meeting Houston (2008) 11. Gonzalez F, Rodrigez M, Galiasso Tailleur R. “Simulación de la alquilación con catalizador solido usando un reactor slurry”. Preprint of the XX SICAT Congress (Brasil), October 17th, (2006), and “Simulation of Three-Phase Spouted Bed Reactor for Solid Alkylation”. (2009) Aiche meeting preprint 581f, San Francisco (2006) 12. Galiasso T. R.; “Simulation of three-phase spouted bed reactor for solid catalyst alkylation” Chem. Eng. Proc. 47 (8), (2008)1384- 1397. 13. Galiasso Tailleur R., Rodríguez S., Farina C. and Derjani-Bayeh S.; “Solid alkylation of C2-C3 olefins with isobutane in the presence of hydrogen using a slurry transport reactorhydrocyclone-regenerator system and PtSO4TiZr/SiO2 catalyst. II Spent catalyst regneration in pilot plant and simulation of FBR

system. Energy and Fuels

XX(cxx)(2017); xxx-xxx 14. Parulekar S., Shah Y.T. Steady State Thermal behavior of gas liquid solid fludized bed reactor Chem. Eng. J. 23 (1982),15-40 15. Sie S.T., Krishna R. “Design and scale-up of the Fischer-Tropsch bubble column slurry reactor. Fuel Processing Technology 64, (2000) 3–105 16. Galiasso Tailleur R., Andretti Salva G. “Cold flow studies of a slurry transport reactor– hydrocyclon system”. Chem Eng Sci. 14, (2008) 3775-3787 17. Galiasso Tailleur R. and Andretti Salva V. “Apparent kinetic rate expressions of light olefins alkylation of isobutane using a PtZrTiSOx/SiO2 catalyst” sent to Catalysis letter (2017). 18. Rough P., Ronaldi J., Galiasso-Tailleur R. “Effect of coke and pore structure in effective Diffusion coefficient of isobutane, 2,3 dimethyl-butane, Hexane, Pyridine and hexylcyclohexane in binary and ternary blends of adsorbates. Spent PtZrTiSOx and Silicalite catalysts. TA&M report 2006. To be publish in Fuels (2017). 19. Ramachandran P., Chaudhari, R.; “Topics in Chemical Engineering” that report Akita and Yoshida (1973), Suganuma and Yamanishi (1966) Hikita and Kikukawa (1975),


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Kato et. al. (1972) correlations, Gordon and Breach Science Publishers, U.S.A., (1983), 308-332 20. Yoo K., Burckle E. C., Smirniotis P. G., “Comparison of protonated zeolites with various dimensionalities for the liquid phase alkylation of i-butane with 2-butene”. Catalysis Letters 74, (2001), 85-90 21. Weitkamp, J., Traa, Y., “Isobutane/butene alkylation on solid catalysts. Where do we stand?; Catal, Today 49, (1999) 193-199 22. Kapteijn, F., Moulijn, J.A., Krishna, R.; “The Generalized Maxwell-Stefan Model for Diffusion in Zeolites: “Sorbate Molecules with Different Saturation Loadings”. Chem. Eng. Sci. 55(15), (2000) 2923-2930 23. Xiao J. and Wei J.; “Diffusion mechanism of hydrocarbon in zeolites”. MIT Thesis (1990) http://hdl.handle.net/1721.1/33810.; Chem. Eng. Sci. 47(5), (1992) 1143–1159 24. Carberry, J, “Chemical and Catalytic Reaction Engineering”, McGraw-Hill, U.S.A., (1976), pages 22-24; 95; 358; 556-558; 580-583 25. Chao, K. and Seader, J.; “A general correlation of Vapor-Liquid Equilibria in Hydrocarbon Mixtures”; AICHE J. 7 (4), (1961) 598-605 26. Galiasso Tailleur R.; “Hydrocyclon-settler-striper design”; Preprint Aiche meeting Houston 2013, sent to Chemical Engineering J. (2017). 27. Svarovsky, L., “Solid-liquid Separation”, London-UK:Butterworths, (1981) pages 222245 28. Rodriguez Dominguez S. “Simulation de un reactor slurry y un regenerador”. Bs Thesis USB, Venezuela (2004) 29. Pro II. SimSci-Esscor's PRO/II software; release July (2008) 30. Turton R. Richard C, R.C., Wallace B. Whiting W. B., Joseph A. Shaeiwitz J.A., Bhattacharyya D., Analysis Synthesis and design Chemical Process (4th Ed.) Prentice Hall International series in the Physical and Chemical Engineering Sciences) 4th Edition ISBN-10: 0132618125 (2008) 31. Galiasso Tailleur R.; “Regeneration of PtSOxTiZr/SiO2 by Hydrogen treatment” I and II, Technical Report Texas A&M Colleges Station (2005-2006). Presented at New Orleans Aiche Annual Meeting, March (2017). In preparation for Applied Catalysis Journal (2017).


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32. Luo X., Lee D.J., Lau R., Yang G., and Fan L.Sh.; “Maximum Stable Bubble Size and Gas Holdup in High-Pressure Slurry Bubble Columns” Aiche J. 45(4), (1999), 665-668 7. NOMENCLATURE a2 a3 a4

Activity, riser inlet and outlet, after downcomer

[-]

aeff

[m2/m3]

A

Effective area of mass transfer slurry-solid ~equal to external area of solid (STR) Constant in adsorption term

[-]

a s, a p

Superficial and particle areas

m2

B

Constant in equation 53

[L/min]

Cpi

Heat capacity of compounds i: gas liquid and solid

[J/mol.K]

Cµ

Drag coefficient defined by eq 57

[-]

D, Dr, DD

Reactor diameter, riser and downcomer diameter (STR)

[m]

Dc

HCS diameter

[m]

Di, Deff

Molecular diffusivity, effective diffusivity

[cm2/s]

DSL

Dispersion coefficient

[m/s]

dp

Particle diameter

[microns]

Dr

Tubes diameter (heat exchanger STR)

[m]

E

Equations (54)

[L/min]

Ei

Activation energy

[kJ/mol]

F

Constant in equation (62)

[Cal/K]

Fi

Molar flux of compound i

[mol/s]

Fi o

Molar flux fed of compound i

[mol/s]

Fw

Mass flow rate of solid

[mT/h]

G

Gravity

[N/kg]

G

Equation (65)

[cal/K]

hw

Heat transfer FBR

[N m2K]

Keq

Equilibrium constant

[-]

KpL

Partition coefficient

[-]

Kp

Adsorption equilibrium constant

[-]

Ki

Preexp. kinetic rate constant

[(mol/l)n-1/i/s ]

kGL,kL, kL.S

mass transfer coefficient at gas liquid and interface

[m/s]


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L

Reactor length

[m]

FH2

Moles flow of hydrogen

[mol/s]

Falk

Mole flow rate alkylate

[mol/m2s]

Nu

Nusselt number for slurry (STR)

[-]

P

Pressure

[Pa]

Pe

Peclet number

[-]

Pi

Partial pressure of compound i

[Pa]

Pr

Prandtl number

[-]

Qi

Fraction desorbed

[-]

R

Radial parameter

[m]

R1,R2

Recycle ratio liquid and solid (PFRR model) Eqs. (56,57)

[mol/mol]

R

Universal constant gas (8,314)

[J/mol.K]

Re

Reynolds number

[-]

ri

Rate of reaction i

[mol/g.s]

S

Reactor section

[m2]

Sc

Smith number

[-]

Tg

Gas residence time

[h]

TL

Liquid residence time

[h]

Ts

Solid residence time

[h]

To, Tf, Tav

Temperature at inlet, outlet and average

[K]

U

Heat transfer in the heat exchanger

ug,fb,

Gas linear velocity disengage zone in STR and in FBR)

[m/s]

vo, vt vmf

[m/s]

Vi

Gas velocity for bubbles, minimum fluidization, terminal gas velocity (FBR). Linear velocity slurry and bubble (riser), slurry in downcomer (STR) Volume of zone i STR and FBR

Z

Axial coordinate

[m]

wo wi

Mass of catalyst

[m t]

xi,

Olefins conversions

[-]

USL,Ug,USLD

Symbols


w/m2 K

[m/s] [m3]

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∆Hr

Reaction enthalpy

[J/mol]

∆T

Temperature difference

[K]

ϛs

Porosity of the solids

[-]

εg εl

Hold up, gas and liquid

[-]

σ

Surface tension interphases

[N/m]

ηi

Effectiveness factor for i reaction

[-]

λeff:

Thermal conductivity of the slurry

[W/mK]

ρsL

Slurry density

[g/m3]

µG: µLS

Gas and slurry viscosity

[cp]

υsL

Slurry superficial velocity

[cm/s]

ρc, ρs, ρl, ρg

Bulk density of the continuous phase, solid, liquid, and gas

[g/m3]

τo

Tortuosity factor

[-]

Captions Tables Table 1: Properties of spent, and fresh plus regenerated catalyst Table 2: Reaction rates expressions Table 3: Kinetics rates constants for the eleven lumps model. Table 4: Mass and heat transfer equations used to simulate the STR Table 5: Catalyst properties for STR simulations. Table 6: Input of data for STR-HCS simulation Table 7: Base case output information for STR-HC at given olefin conversion and inlet temperature Table 8: Effect of temperature and particle size in product composition (other operating conditions of base case). Table 9: Effect of changing feed throughput and temperature respect to the base case STR Table 10: Effect of particle size in reactions rates (Semi-batch reactor). Calculation of STR volume as function of particle size. Table 11 Effect of mass transfer on main reactions


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Captions Figures Figure 1. Process flow scheme. Right side STR and FBR zones; left side Separation and recycling zone Figure 2 2a). Batch and semi-batch reactor; 2b) Alkylation pilot plant (CSTR). Figure 3 3a) Ethylene conversion (Y left axis) and coke production (right Y axis) at different residence times and temperature (points is the average value, bracket represent the extreme value observed measured in repeated experiments); simulation of ethylene conversion and coke production (dash lines); 3b) Effect of particle sizes, temperatures, products of reaction and pressure in conversion of olefins (right Y axis, bars). Effect of catalyst efficiency factor η in selectivity at different particle size (dash violet-lines) and pressure (dash red-lines) Basic operating conditions are: 0.2: O/IC4, gas/liquid: 0.3, 100 rpm, PT:1.4 MPa, PH2:0.88 MPa; cat/HC:0.08, values of conversion and selectivity are calculated by extrapolation to t=0. Figure 4. Pilot plant long-term test. Left axis conversion of olefins into ● alkylate as a function of time on stream (C2+C3+IC4→). Right axis Conversion of olefins into ● coke and ● paraffins 343 K; 1.4 MPa; O/IBO: 0.3, H2/O: 5 (moles), cat/HC: 3.8 % wt; particle size variable. Figure 5. Carbon distribution in product for 40 to 60% of conversion range. Figure 6. Predicted by the simulation model vs. measured values in the long-term pilot plant test: (Left Y axis) Olefins to alkylate; (middle Y axis) olefins to coke; (Right Yside) olefins to paraffins. Values measured at different operating temperatures and gas, liquid and solid flowrates (1.4 MPa; O/IBO: 0.3, H2/O: 5 (moles)). Figure 7. Left side Y-axis: delta Carbon predicted by the model vs. delta coke measured in the pilot plant for three levels of carbon on inlet catalyst. Samples obtained at different temperatures, gas, liquid and solid flowrates, and particle diameters. Right side. Startup of pilot plant. Activity of catalyst as a function of time on stream. Steady state value of activity predicted vs measured at different operating conditions Figure 8. Details of the alkylation (STR), separation (HCS) and regeneration (FBR) zones Figure 9. Nomenclature for mass and heat balance in STR and HCS Figure10. Program used for simulating the STR. Visual basic 6. Output: report of concentration and temperature profiles, components flowrates (m t/h) and cost of alkylate $/ (m3 octane) Figure 11. Example of results of simulation. Concentration profile of olefins isobutene and hydrogen as a function of riser reactor volume (STR) for the base case operating conditions (373 K, H2/olefin: 2 (moles), O/IB: 0.3, 1.4 MPa, 4–60 microns, TL: 0.4 h and Tg: 0.3 h. Figure 12. Results of simulation. Right Y-axis activity vs. reactor volume at three inlet temperatures, Left axis delta of temperature along the riser and downcomer reactor as a function of volume at three inlet temperatures. Other operating conditions at base case: H2/olefin: 2, O/IB: 0.3, 1.4 MPa cat: 40-60 microns, TL: 0.31 h and Tg: 0.2 h. Figure 13. Effect of gas-liquid velocities (T: 363 K, dash lines) and inlet temperatures (ug: 0.25 m/s; full lines) in recycle ratio (left Y-axis), number of stirred tank in series


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Page 50 of 68

(right upper Y-axis) and in bubble diameter size (right lower Y-axis). Other operating conditions at base case: H2/O: 2, O/IC4: 0.3, 1.4 MPa Cat/HC: 0.08; Cat: 40-60 microns, TL: 0.3 h).

Annex I. Complementary information Extensive experimental studies have been performed in reactor operating in batch mode (V: 0.5 L) with the above-mentioned catalyst and using different concentrations of ethylene, propylene and isobutane; the molar ratios were combinatorial of either: 0.05 - 0.07- 0.1 for olefins/isobutane (O/IC4), 0.2 - 0.3 for hydrogen /olefins (H2/O), 0.5 -1 - 2 for C2/C3. For kinetic rate study, the catalyst particles were initially diluted in either IC4, 0.9 IC4 + 0.1 of 2,3 DMB, 0.9 IC4 + 0.1 of 2,3 TMB or 0.9 IC4 + 0.1 of HCH and stirred at 100 rpm until the reaction temperature is reached. At t=0, olefins are added, and the reaction started. After 0.1 hours of contact time (Run 1) gas-liquid-solid were withdraw from the reactor and samples analyzed. The same is performed for Run 2, 3, 4 and 5 where contact time were set to 0.2, 0.3, 0.4 and 0.5 hours. Each of tests were repeated three times. Kinetic rates equations were confirmed using a total 56 experiments based on 53 experimental design. Apparent kinetic rates model, previously developed for butane alkylation9, was seeded to the program to simulate olefins disappearance and product yields at different operating conditions and data obtained were compared to experimental values. Kinetic rates constants were obtained using a computerized program based on Genetic algorithm (optimization tool, see below) that adjust pre-exponential and activation energies constants until less than 0.1 % difference in produces yields respect to experimental are predicted. Concentration of gas liquid and solid were calculated with the following equations:

(

γ y V . δ −δ (0 ) (1) 0 K i ≡ i = ν 0 i i (91); logν = logν + ω logν (92), lnγ i = i i R.T xi φi

)

2

(93), δ =

∑x .V .δ ∑x .V i

i

i

i

i

(94)

i

i

logν (0 ) = A0 +

(

)

A1 2 3 2 2 + A2Tr + A3Tr + A4Tr + A5 + A6Tr + A7Tr Pr + ( A8 + A9Tr )Pr − log Pr Tr

(95) logν (1) = −4.23893 + 8.65808Tr −

1.22060 3 − 3.15224Tr − 0.025(Pr − 0.6 ) Tr

(96) lnφi = ( z − 1)

Bi A2  Ai Bi   B.P  − ln( z − B.P) − 2 − ln1 +  B B  A B   z 

(97) On surface concentration were obtained by simulation at different operating conditions the gas liquid and solid equilibrium and then data were regression to obtain a polynomial expression of concentrations as a function of operating variables (Ci, T and P). See detail in references12,13. These equations are included in the simulation program. Batch reactor pseudo isothermal mass balance equations as function of contact time (t) are: Gases


Page 51 of 68 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

Ethylene: KGLK p,C2[a(t)η1 (k 1CC2 CCIC4 + k3CC1.25,l (C6,L + C7,L ) + k5a(t)η1CC2,L (CC8,L + C9,L ) + a(t)k9η1CC2,LCH0.32,L )] dCC2 = KGL(CCz 2.g − CCf 2,L ) − dt [ A + KIC4CIC4 + KC6−10C6−10 ] − [k7CC1.25,LCC8−C10,L / 1 + 4CH 2,L )

(98) Propylene: K GL KpC 3 a (t )η 2 (t ){k 2CC3 ,LCCIC 4 ,L + k 4CC 3,L (CC 6 ,L + CC 7 ,L ) + k 6CC 3,L (CC 8,L + CC 9 ,L ) + k10CC 3,LCH0.32,L } dCC 3 = K GLCCZ3. g − CCf 2 ,L ) dt [ A + K IC 4C IC 4 + K C 6−10C6−10 ]

(99)

− k8CC1.35,L ∑ CC 8−C10,L / 1 + 4CH 2,l

Isobutene: K K a(t )η1 (t )(k1CC 2,L + η 2 k2CC 3, L )CCIC 4,L d IC 4 = K GL (CiCz 4. g − C ICf 4, L ) GL p ,IC 4 dt [ A + K IC 4, LC IC 4,L + K C 6−10,L CC 6−10,L ]

(100)

Alkylates (C6-7): dCC 6−7 = K GL (−CCz 6−7 ,L + CCf 6−7 ,L ) − dt

101)

K GL K p ,C 6−7{a (t )(qk11[(η1 (t )k1CC 2,L + η 2 (t )k 2CC 3,L )]CCIB , L } − a(t )(1 − q )k11{η1 (t )[k3CC 2,L (CC 6,L + CC 7, L ) + η 2 (t )C3,L [(k5 (Cc6,l + Cc7 ,L )]} [ A + K IC 4CIC 4,L + K C 6−10CC 6−10 , l ]

Alkylate (C8-10): dCC 8−10 = KGL (−CCz 8−10, g + CCf 8−10, gL ) dt

(102)

KGL K p,C8−10{a(t )qk11[(η1 (t )k3CC 2,L (CC 6,L + CC 7 ) + η2 (t )[k5CC 3,L (CC 6,L + CC 7,L )]CCIB,L − (1 − q)k11 (η1 (t )k7CC1.25,LCC8−C10, L − η2 (t )k8CC1.35,LCC8−C10,L ) /(1 + 4CH 2,L } [ A + K IC4CIC 4, L + KC 6−10CC 6−10,L ]

Light paraffins: [a (t )(k9CC 2, L + k10CC 3, L )C H0.32, L )] dCC 2+C 3 = K GL CCz 2. L − CCf 2,L ) − dt [ A + K IC 4C IC 4,L + K C 6−10CC 6−10, L ]

(103)

Soluble Coke: dCC10 + , s (1 − q)k11K GL K pC 8 −10 [η1 (t )k7CC1.25, LCC 8 −C10, L /(1 + 4CH 2, L ) + η 2 (t )k8CC1.35, LCC 8 −C10, L /(1 + 4CH 2,l )] = [ A + K IC 4CIC 4 , L + KC 6 −10CC 6 −10, L ] dt

(104)

− k12CC10 + , s

dC C10 + (105) = k12 CC10 + , s dt Notice that the equilibrium constants that convert on-surface [Ci] into liquid-concentration Ci are dependent of partition coefficient K oi , LSeq (liquid solid equilibrium of adsorption),

Insoluble coke:

Koi ,GLeq gas-liquid equilibrium constant and VSL (molar volume of the slurry). All of them are 0

integrated into the pre-exponential factor ki depicted in Table 2 and 3 and the heat of adsorption and heat of gas-liquid equilibrium are integrated in the apparent activation energy Ei. ηi can be simplified by using the following assumption: Deff , DMB ~ Deff ,C 2 ~ Deff ,C 3 : 0.29 E − 0.8cm 2 / s (106); activity is proportional to the open intersection micro-mesopores or the amount [C10+] (coke) adsorbed there and inversely proportional to hydrogen partial pressure.


Energy & Fuels 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 52 of 68

Similar equation kinetic rate expressions are used in mass and energy balance equation developed for pilot plant CSTR, operating in transient or continuous steady sate mode. A fluid-dynamic model was used to take in account bubbles going up and catalyst additionwithdrawal. CSTR is connected to a gas-liquid-solid separator. The reactor has a constant slurry- gas phase interphase and it is operated at constant temperature. The reactants get dissolved in the slurry phase, then adsorb in the solid and react. Bubbles goes up in a piston flow movement exchanging reactant and product with the slurry. Slurry is a perfectly mixed phase. Overall transient mass balance of hydrogen can be written for example as: For gas phase: V g d ( PH 2 / RT ) dt For liquid V L d (C H 2 , L ) dt

=

v gf .PH 2 RT

−

v go .PH 2 RT

− k GL a (

PH 2 , RT

= v Lf .C H 2, f − v Lo .C H 2 , 0 + k GL aGL (

−

PH 2 , RT

C H 2, L KG−L −

(108)

)

C H 2, L K G−L

) − k LS a LS (C H 2 L − C H 2 , S )

(109)

For d (Vs C H 2, S )

solid

dt

= k LS a LS (C H 2 L − C H 2, S ) − Vs ∑ aη i ri

(110) In the solution of these equations, we have not considered the accumulation in solid, gas and liquid because the system interphase is automatically control LIC during the transient and solid coming in and out is controlled by an automatic feeding and withdraw algorithm. Integration of time derivatives was carried out using LSODA routine which can handle stiff or non-stiff ordinary differential equations. For kinetic rates analyses, these equations were simplified and only the steady state solution (CSTR) were used. Transient analysis was used to evaluate the effect of solid flow rate need in the approach to steady state (Fig. 7b). Tables Table 1

Spent D33

Properties Pt wt% ZrO2 wt% bulk /surface ITiO2/ΣIMe% TiO2 wt% bulk / surface IZrO2/ΣIMe% SO4= wt% SiO2 wt% Surface area m2/g Particle 10-40microns diameter % Pore volume cm3/g Micropores volume cm3/g Tortuosity factor* Benzene Diffusivity cm2/s 373 K (thermo-balance Benzene/Ar)

95.5% regenerated D43 + 4.5% fresh 0.15 3.2 / 4.2 3.1 / 3.9 5.4 5.1 complement 108 190 0.2 0.1 0.29 0.39 0.08 0.08 0.86 0.98 1.3x10-5 3.2x10-5


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Energy & Fuels

Particle size microns Pyridine remain ads. 373 K µmol/g Carbon content % wt. Catalyst activity for alkylation** Coke soluble in CS2 (323 K) Naphthenic compound in coke % % of C12 in CS2-extract

40-60 0.14 0.81 0.23 0.62 70 45

0.28 0.13 0.83 0 11 0

*Measured by nitrogen diffusion in a Wicke-Kallenback type of cells** measured in semi-batch reactor at t=0 (363 K. 1.4 MPa)

Table 2

 Eo  w − r1 = k1'o . exp − 1 .CC2= [CCIc4,S ]a wo  R.T  o  E  w − r3 = k 3'o . exp − 3 .C C2 = .[CC6, s + CC 7 , s ]a wo  R.T   Eo  w − r5 = k5'o . exp − 5 .CC2= .[CC8,s + CC 9, s + CC10, s ]a R . T w   o o  E  w − r2 = k 2'o . exp  − 2 .C C3 = .[C IC4, S ]a wo  R.T 

 Eo  w − r4 = k4o. exp − 4 .CC3=.[CC7,S + CC6, s ]a wo  R.T   Eo  w − r6 = k6'o.exp − 6 .CC3=[CC7,S + CC8,s + CC10,s ]a wo  R.T   Eo  k7'o .exp − 7 .CC1.52 = ∑[CC10−C13 ] w  R.T  − r7 = a. 1 + 4.CH2.,S wo

 Eo  k8'o . exp − 8 .CC1.35= ∑[CC10−C13 ] w  R.T  − r8 = a. 1 + 4.CH2,.S wo

(9)

(10)

(11)

(12)

(13)

(14)

(15)

(16)

 Eo  w − r9 = k 9'o . exp − 9 .CC2 = .[C H0.52, s ]a wo  R.T 

(17)

 Eo  w − r10 = k10'o . exp − 10 .CC3 = .CH0.52 , s a wo  R.T 

(18)


Energy & Fuels 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 54 of 68

 Eo  w − r11 = q p k11o . exp − 11 .. ;  R.T  wo

(19)

qp = exp(−Ci / 16)

(20)

Table 3

Reaction Alkylation ri (mol/g h) 1 2 3 4 5 6 7 8 9 10 11 qIC4/qC6-C7/ qC8-C10/ qC10+

 KJ  E   mol 

75.3 82.2 70.6 71.2 122.0 73.1 72.2 132.0 92.4 87.8 64.8


  Ln −1  k io  n −1  Kmol min 

6.00E+09 8.00E+10 2.50E+08 2.00E+08 4.00E+14 2.50E+08 3.00E+08 1.40E+16 2.70E+10 5.60E+09 5.90E+07 1/0.8/05/0

Page 55 of 68 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

Table 4 dX 1 − r1 .ε L = = f1 ( X i , T ) dV F ° Ethylen -Liq

(21)

− r3 .ε L dX 3 = = f3 ( X i ,T ) dV F ° Ethylen - Liq

− r2 .ε L dX 2 = = f2 ( X i ,T ) dV F °Ethylen -Liq

(22)


(24)

(23) − r5 .ε L dX 5 = = f5 ( X i ,T ) dV F °Ethylen-Liq


(25)

(26) − r8 .ε L dX 8 = = f8 ( X i , T ) dV F ° Ethylen -Liq

− r7 .ε L dX 7 = = f7 ( X i ,T ) dV F ° Ethylen -Liq

(28)

(27)

− r9 .ε L dX 9 = = f9 ( X i ,T ) dV F °Ethylen - Liq

− r10 .ε L dX 10 = = f10 ( X i , T ) dV F ° Ethylen -Liq

(29)

(30) FHydrogen,55 = FHydrogen,54 − 3 X r ⋅ FCoke,27

− r11.ε L dX 11 = = f11 ( X i , T ) dV F ° Ethylen - Liq

(32)

(31) FC3,55 = 0.22 X r ⋅ FCoke,27

F 'Ethylene = FEthylen,7.(1 − X1 − X 3 − 2 X 5 − X 9 )

(34)

(33) FC4,55 = 0.44 X r ⋅ FCoke,27

FC5,55 = 0 .22 X r ⋅ FCoke,27

(36)

F 'Hydrogen = FHydrogen , 7 − ( X 9 + X 10 ) ⋅ FEthylen,7

(38)

(35)

F'Propylene= FPropylen,7−( X2 + X4 + X7 +2X8 + X10)⋅ FEthylen,7 (37) FCoke,28 = FCoke,27 (1 − X r )

(39)

F 'Propane = FPropane,7 + X 9 ⋅ FEthylen,7

(40)

F 'Ethane = FEthane,7 + X 5 ⋅ FEthane,7

(41)

F 'Alkylate1= FAlkilate1,7 + ( X1 + X 3 + X 6 )FEthylen,7

(42)

F 'Coke = FCoke,11 + ( X 5 + X 8 ) ⋅ FEthylen,7

(44)

F 'isoButane = FisoButane,7 − ( X1 + X 2 ).FEthylen,7 (43)


Energy & Fuels 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 56 of 68

F 'Alkylate2 = FAlkylate2,7 + ( X 2 + X 4 + X 7).FEthylen,7

F 'Propane = FPropane,11 + ( X10 ) ⋅ FEthylen,7

(45)

(46)

F 'Ethane = FEthane,11+ ( X 9 ) ⋅ FEthylen,7

τs =

(47) εP =

LF ° Catalizado r PM P Vu TP ρ P

(ao wo + a1w1 − af ) kd (1.8 − kd C10+, f )

εP =

(73)

(48)

F ° Catalizado r (1 − ε G ) υ L ρ P + F º Catalizado r

(74)  g ⋅ (ρ P − ρ L )  U SL = 0.0178 ⋅  ρL ⋅ µL   2

2

1

3

⋅ dP

(75)

εg 1.6(uo ρ g )0.3 ( ρ sl − ρ g / ρ sl )0.23 = 2.33 (1 − ε g ) (0.4ξµ ) d p0.33[cosh( 3 l g )0.04 ]2.3 σ ρ sl db =

(77)

w

=

wo

4 2.98σ g µ SL ( ) 0.25 3 u b ρ SL ρ Lσ

ug dT

(81)

DSL

k \ L ,S .d P Di

(u

TP

.L DEP ).e− uTP . x DEP 1 − e −uTP . L DEP

ln µ sl = ln µ L + ln 1.2

(79) 2 C µ = 0.22 − 0.263 ⋅ log(µ SL ) + 0.103 ⋅ (log µ SL ) R2 =

− ρg 2.1σ 0.85 ρ + 0.2 gd p ( SL ) ρ sl d p ρ SL

ub =

0.3 < uo < 1.0m / s

d p / DEL JS +1 Js = 0.25(W / Wo ) uo L / DEl 14

 u g.d 4 ρ 3  = 2 + 0,4. G. P3 L  µL  

 µL   ρ L .Di

13

  

(83)

;

(85)

(78)

w gµ L4 0.23 w [exp − ( ) − 1.4 ] wo ρ Lσ L wo

−0.85 −0.85   ug    ug   1+ 8   / 1+ 8     gdT     gdT        0 . 02 gD d / D c + 6 . 6U 1g.25 ρ s− 0 .5

 ug = 13  gd T 

(N ) =

(U g )

ρ L µL ε g σ L 0.3U g 0.2d s 0.15 0.4

KG −L = 0.7 x104

0.2

1.4

(86)  u G . g .( ρ SL − ρ g ) wsZ '= 0 (%) =  3  µL σ 

3

(87)

0.3

  C D0.5  

−0.14 wsZ = Z = 1.23e wsZ = 0

µ g d 2p ( ρ s − ρ g ) Z g

(88)

Table 5 0.5% Fresh + 99.5 % D43 Regen. catalyst 40-60

Catalyst Properties Diameter,

dp ; µm

Spent catalyst D33 1-60

Density, ρ p ; gr cm3

1.19

1.44

Catalyst average activity

0.81

0.19

Molecular weight, MW; gr mol

106

118

Specific Heat, C ; J P

139+32.10-3T

grF

Average micro-pore diameter Ǻ (BET) Effective diffusivity Deff,DBM-IC4 (cm2/s) 373 K liquid phase23 Effective Conductivity, K eff ;

J m sK 2

5

8

0.29x10-9

0.21x10-9

5.77 10-6

8.8 10-6


(76)

(80)

(82) (84)

Page 57 of 68 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

Table 6

Input Data Inlet feed Temperature (K) C2- C3, iC4 Inlet solid Temperature K Olefin flow [≡ ]

m3 D

Set Value 363 543 783

Pressure MPa Ratio fresh /regenerate catalyst wt. Ethylene + Propylene conversion (-) H2 to olefin ratio (-) Olefin to IC4 ratio (-) Alkylate to olefin ratio (-)

1.2 0.02 0.52 2.0 0.1 3.0

Mass of catalyst to HC wt. %

8.0

ρg (kg/m3) (363 K, 1.4 MPa)

76

KGL cm2 sec-1

2.8*10-2

KL cm2 sec-1

0.8*10-2

σL (N/m) (363 K, 1.4 MPa)

0.33

ρL (kg/m3) (363, 1.4 MPa)

728

µL (Pa.s) (363 K, 1.4 MPa

0.051

uo,g (m/s) riser

0.12-0.2

uo.l (m/s) riser

0.01-0.015

U (W m-2K-1) bundle of vertical tubes D/L ratio slurry zone, STR (-)

450 7

DHCS1/ DHCS2 cylinder (m)

0.4 / 0.2

L 1, HCS1/ L 1, HCS2 cylinder (m)

0.5 /0.3

L2, HCS1/ L 2, HCS2 cone (m) Gas liquid solid top efficiency HCS (-)


0.4 / 0.25 0.99/0.97/0.02

Energy & Fuels 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 58 of 68

Table 7

Simulation Results (BC) Outlet Temperature K 3

STR Volume m of reaction zone (V1+V2)

Value 396.3 43+14

Ethylene conversion into Alkylate 1 (-)

0.48

Propylene conversion into Alkylate 2 (-)

0.63

Ethylene conversion into Coke (-)

0.02

Propylene conversion into Coke (-)

0.01

Ethylene hydrogenation to ethane (-)

0.014

Propylene hydrogenation to propane (-)

0.018

Gas holdup (-)

0.10

Average Bubble size (m)

0.08

Liquid holdup (-)

0.81

Solid holdup (-)

0.09

Bubble velocity (m/s)

0.22

Internal recycle (R1)

0.35

Liquid residence time hr

0.31

Catalyst activity at the outlet (a2) (-)

0.25

Carbon on catalyst wt%

0.84


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Energy & Fuels

Table 8

Temperature K / % wt Ethane Propane Butane Pentane Methyl-butane Cyclopentane Hexane Cyclohexane Methyl-cyclopentane 2-Methyl-pentane 2-Ethyl-butane 2,3-Dimethyl-butane 2,2-Dimethyl-butane 2 Ethyl 3 methyl-butane Methyl-cyclohexane 2,2-Tri-methyl-butane 2,3Ethyl-pentane 2 Methyl-hexane Propyl-butane 2,3 ethyl methyl but. 2,2,3 Trimethyl-pent. Ethyl-cyclohexane

375/50 0.002 0.005 0.005 0.005 2.167 0.025 0.837 1.133 0.049 10.345 2.562 16.158 15.123 23.300 0 9.901 6.355 3.990 1.059 0.567 2.217

380/50 0.002 0.006 0.007 0.010 3.333 0.049 0.637 1.250 0.044 11.029 2.500 14.716 13.578 22.157 0.01 10.343 6.520 4.510 0.907 1.225 2.206

385/50 375/150 0.005 0.002 0.008 0.005 0.010 0.005 0.012 0.005 3.812 2.022 0.099 0.033 0.495 0.827 1.287 1.258 0.050 0.052 11.881 9.865 2.475 2.662 12.861 15.231 13.465 15.011 20.792 22.600 0.02 1.22 11.015 9.962 6.683 6.389 5.099 3.073 0.619 1.042 1.733 0.593 2.178 2.201

0

0.01

0.03

0.81

C8-C10

4.286

5.000

5.842

4.301

Octanes (experimental) (-)

98

99

99.5

99

Liquid yields V%

94

92

91

92

Table 9

Alkylate 1

Alkylate 2

Coke

Paraffins

$/(m3 xOctane)

TL, BC(h), TBC

19.23

28.72

0.07

0.14

25.0

TL, BC+ 20% (h), TBC

17.45

26.45

0.05

0.19

27.2

TL, BC- 20% (h), TBC

20.60

30.44

0.07

0.19

24.6

TBC+ 8K, TL, BC

19.93

37.69

0.24

0.18

26.3

TBC+ 8K, TL, BC

18.29

27.99

0.05

0.14

24.8

m t/h


Energy & Fuels 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 60 of 68

Table 10

Parameters Base case operating conditions, 40-60µm cat. X1+X3+X5, Ethylene conver. to alkylate 1 (wt) η1:0.88 X2+X4+X6, Propylene Conv. to alkylate 2 (wt) η2:0.82 X7, Ethylene conversion to coke (wt) X8, Propylene conversion to coke (wt) X9, Ethylene conversion to ethane (wt) η5 X10, Propylene conversion to propane (wt) η6 Volume ( m 3 ) for 65% Olefins conversion. Alkylate yields m t/h T° (K) inlet TF (K) outlet Final catalyst activity Coke on catalyst

Without mass transfer

With Product & reactant mass transfer control

0.632

0.512

0.859

0.532

0.006

0.025

0.005

0.02

0.002

0.0018

0.002

0.0018

17.2

58

65 373.0 406.34 0.45 0.55

49 373.3 402.1 0.24 0.83

Table 11

Rates gmol/h m3 at 50% conv.

dp = 50 µ m

dp = 100 µ m

dp = 150 µ m

k A,L .ai .C A,L gmol/L h

13.0

11.7

10.2

k A,S .aC .CA,L gmol/L h

5.9 0.7

5.3 0.59

4.8 0.48

0.04 57 0.81

0.06 89 1.1

0.10 167 1.54

∑ η a.(− r ) gmol/L h i

(r

* 7

i

i

)

+ r8* gmol/L h

*Reactor volume m Coke wt %

3

*set to produce 49 m t/h of alkylate

Figures


Page 61 of 68

Figure1 Simulated in STR and FBR programs

H 2 purge

V-106 22

F-104

10 E-101

13

12

14

E-103

V-102 15

M-102

1

STR-101

E-105

E-106

7

T-103

lps

28

47

46 V-111

E-104 18 42

F-101 M-101

T K-101

4

F-102 P-101 A/B

53

5

iC4,nC4

Alkylate

E-109 50

cw

52

From main fractionation

51 E-113

E-112

P-103 A/B

Figure 2 Solid make up in slurry Flow stirring and pressure control H2 Olefins +IC4

∆P Bubble Frequency Analyzer

CSTR (5 L) Heat control

Gas To GCMS

FRC HPLTS

Reaction products

liquid To GCMS

Olefins +IC4

vacuum

Solid To CNMR and C analyses.

a)

Gas To GCMS

HPLTS

H2

liquid To GCMS CS2, Ar

H2 FRC

Heat control

Semibatch or CSTR 0.1 L

b)

Figure 3 Batch reactor data 0.1

coke

0.9

0.2

0.8

0.3 0.4

0.7 0.6

Run 1

0.5

0.5 Run 2

0.6 0.7

0.4 0.3

Run 3

0.8

0.2

Run 4

0.9

0.1 0.1

0.2

0.3 0.4 Contact time (h)

0.5


Conversion ethylene into Coke wt%

T K-102 i-C4,C4

V-112 38

39

P-102 A/B

22

20

E-110 lps

40 6

2

19

41

Cat withdrawal

C-101 A/B

iC4,nC4

De-butanizer

H2

54

,C3, ,C3=

44 45

55

49 V-110

43

17

27

FBR-102

V-103

30

48

E-111

V-109

26

Cat make up

TK-103

,C3+ C3=

35

De-propanizer

31

11 25

V-108 E-108

lps

V-107

37

36

E-107

34 16

HCS-102

C-102 A/B

Olefins feed

33

8 HCS-101

F-103

To fractionation

32 T -102

E-102

C-103 A/B

H2 Make up

H2,C2,C3,C2=,C3=

V-101

9

3

to main fractionator

29

24

V-104

Simulated in PRO II

23 V-105

Conversion ethylene into C 6 (-)

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

b)

50µm, 353 K, M 0.8 MPa

0.85 50µm, 353 K, 1.8 MPa

50µm,353 K, 0.1 M DMB

0.9

50µm,353 K, 0.1 M DMP

50µm,353 K, 01 M HCH

50µm,373 K 100µm, 373K 150µm, 373 K

0.2

150µm, 363 K

0.4

50µm,363 K 100µm, 363K

0.6

50µm,353 K 100µm, 353K 150µm, 353 K

Olefins conversion (-)

Run 3

0.8

Product of Pressure Alkylation

Particle size and Temp

Figure 4

Product distribution 30

0.3

25 20

0.25

C10-16

C7

0.2

Nap/Par

C6

15

0.15 Sol. Coke

10

0.1 C9 C8 C4+C5 C2+C3

5 0 35

45 55 Conversion of Olefins wt%


0.05 0 65

Coke, Nap/Par ratio, C10-16 Hydroc. wt%

Figure 5

C5-C9 Alkylate wt%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 62 of 68

(Selectivity to alkylate (-)

Energy & Fuels

Page 63 of 68

Figure 6

60

O→ALK

O→Par

O→C10+

55

0.4

0.4

50

0.3

0.3

45

0.2

0.2

45

50

55

60 0.2 0.3 0.4 0.5 0.2 Measured Conversions wt%

0.3

Figure 7 dp: 175 microns dp: 100 microns dp: 50 microns

0.8

FBR Pilot plant

0.6

0.4 90

95

100

dp: 175 µm dp: 100 µm dp: 50 µm

0.8

Predicted a (-)

Predicted ∆a (-)

Predictd Conversions C2 =+C3= wt%

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

Steady state

0.6

Pilot plant Steady State

0.4 0.2

Pilot Plant: Start up 50 microns, 360 K

transient state a)

0.4 90

950.6

10080

Experimental ∆a (-)

0.2

b)


0.4

0.6

0.8

Measured a (-)

0.4

0.5

Energy & Fuels

Figure 8 Gas to separation (PSA) Gas –solid

V3

Liquid to separation Hydrocyclone

V2

Solid Settler V1

Fresh Cat.

STR

Solid Stripper Hydrogen to purification Spent catalyst

V6

FBR

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Page 64 of 68

C2/C3 Temp. control

Hydrogen recycled

V4

Isobutane fresh + recycled

Figure 9 gas (alk)

2 PH2 2 , PHC T2

3 PH3 2 , PHC , T3

V1 , ε gr

0 PH0 2 PHC T0

go

R PHR2 , PHC , T0

5 PH5 2 , PHC , T5

V3

Bubbles PFR

To Fract.

HCS

C HS 2 , T s 1

V 2 R1; recycle

4 PH4 2 PHC PT4

Liquid (alk)

STR

2 2 C C T C HC C H 2T2 1 H2

0 CHC ,T0 vo

v1

V1 , ε L

v2

HCS

Slurry PFR

V 2 R1; recycle

Solid (alk-reg) m1

Regeneration 9 Fsreg a9

m10

4 CH4 2 C HC PT4

Fs2 a 2 T2

Fs T4 a 4

8 C H8 2 , C HC , T8 m8 To Fract.

3 C H3 2 , C HC , T3

1 HC 1

R CHC ,T0

Fso ao

5 PH5 2 , PHC , T5 m6 6 6 PH 2 , PHC , T6

Fs14 a 3

Alk WsAlk ,1 , ε s

Fs3 a3 T3 Slurry PFR

Ws2 , R2, recycle

V3

To FBR

HCS

m12 Fs12 a 3 To FBR


Fs11 a 3

m11

Page 65 of 68 1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 19 20 21 22 23 24 25 26 27 28 29 30 31 32 33 34 35 36 37 38 39 40 41 42 43 44 45 46 47 48 49 50 51 52 53 54 55 56 57 58 59 60

Energy & Fuels

Figure 10 Input Data. Define Vol. or Conv. calculation. Define ratios, recycles and operating conditions Define particle size and distribution. Select L/Dc ratio and Dr/Dc. Define maximum amount of solid in gas (m8). Define boundary conditions and efficiency in gas-solid-liquid separation for two HCS. Introduce V1, guess and V2 gue ss T1,guess and T2,guess or x1,guess x2, guess T1,guess and T2,guess Calculate inlet properties: Gas-liquid-solid mass flowrates, recycles, dimensionless number and phases properties at inlet Solve slurry and gas phase set of differential equations along reactor volume using RKF numerical method point by point and boundary conditions. Calculate [Ci], catalyst ratio (wi /wo ) mass transfer (ki ηi) and catalyst activity (a) along the reactor Calculate conversions and temperature at the outlet. If x1 and x2< |xguess ±0.01| and T1f and T2f < |Tgue ss ±0.1K| calculated V1 and V2 or if V1,f and V2,f

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