Joint Transformation of Methanol and - American Chemical Society

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Joint Transformation of Methanol and n‑Butane into Olefins on an HZSM‑5 Zeolite Catalyst in Reaction−Regeneration Cycles Andrés T. Aguayo, Ana G. Gayubo, Ainara Ateka,* Mónica Gamero, Martin Olazar, and Javier Bilbao Departamento de Ingeniería Química, Universidad del País Vasco, Apartado 644, 48080 Bilbao, Spain ABSTRACT: The joint transformation of methanol and n-butane has been studied under energy-neutral conditions (methanol/ n-butane molar ratio of 3/1) on an HZSM-5 zeolite catalyst in order to determine the reactivation kinetics and optimize the process conditions for maximizing the light olefin production rate. The methodology for determining the reactivation kinetics combines the kinetic models for the main reaction and deactivation (determined in previous studies), as well as the experimental reactivation results obtained for a reaction step (at 500 and 550 °C) subsequent to regeneration by coke combustion with air (at 550 °C, up to 120 min). By simulation of the operation in reaction−regeneration cycles, an optimum average olefin production rate of 22 mol/(gcatalyst h) is obtained for a reaction temperature of 500 °C, space time of 0.37 gcatalyst h mol−1, time on stream of 40 min, and partial reactivation by coke combustion for 15 min.

1. INTRODUCTION Methanol upgrading by catalytic transformation (MTO process) is a route with increasing industrial implementation due to a greater demand for light olefins and automotive fuels from fossil fuels other than oil (coal and natural gas) and from lignocellulosic biomass.1,2 Furthermore, methanol and other biomass-derived oxygenates (ethanol, butanol, glycerol, and flash pyrolysis bio-oil) are good potential refinery feeds, particularly in FCC (fluidized catalytic cracking) and hydrocracking units.3,4 Different studies under standard conditions in refinery units have been reported with encouraging results on the compatibility of cofeeding oxygenates with standard hydrocarbon streams.5−9 The joint transformation of methanol (obtained from sources alternative to oil) and paraffins (secondary products in refineries) for producing olefins and fuels is of great interest from the perspective of sustainability and the intensification of crude oil upgrading processes. The main difficulty of the joint transformation process is the low cracking capacity of paraffins, which requires catalyst performance and conditions to be midway between the optimum ones for the production of olefins from both methanol and paraffins. The pioneering studies on this joint transformation have been carried out using an HZSM-5 zeolite catalyst and by cofeeding methanol with n-butane10,11 and with hexane.12 The authors of this paper have studied the kinetic modeling of individual reactions, i.e., transformation of methanol into olefins13 and the cracking of paraffins,14 and the synergy in the joint transformation.15−18 Thus, they have reported the suitable performance of an HZSM-5 zeolite catalyst (SiO2/ Al2O3 = 30), given that 80% selectivity of C2−C4 olefins was obtained (by mass unit of CH2 equivalent units in the reactant mixture) at 500 °C and for a space time of 0.05 gcatalyst h/ molCH2.17 A significant advantage of the joint transformation of methanol and n-butane is the attainment of an energy-neutral regime (for a methanol/n-butane molar ratio of 3 in the feed), given that an exothermic process (methanol transformation) © 2012 American Chemical Society

and an endothermic process (butane cracking) are integrated in the same reactor.16 Duduković emphasizes the interest of strategies based on coupling exothermic and endothermic reactions for process intensification.19,20 These strategies have been described by Ramaswamy et al.21,22 and involve carrying out exothermic and endothermic reactions in the same reactor or in a multitubular reactor with heat exchange through the wall between the individual reactors. In the joint transformation of methanol and n-butane, the main reaction and coke formation corresponding to both individual reactions are carried out in the same reactor, which gives way to synergies:16 (i) n-butane cracking is enhanced by energy compensation on the acid sites; (ii) the autocatalytic mechanism of olefin formation from methanol (hydrocarbon pool mechanism) requires an initiation period and is enhanced by the rapid formation of olefins in the cracking of nbutane;23−25 (iii) the deactivation, which is very fast in the transformation of methanol, is attenuated by inhibition of polyaromatic formation due to the presence of a higher olefin content in the reaction medium;18 (iv) the formation of methane (associated with catalyst deactivation) is attenuated.17 The aforementioned effect of energy compensation is a significant advantage for the integration in the same reactor of an endothermal reaction with an exothermal. Consequently, the real temperature in the acid sites (on which both reactions occur) is close to the nominal one. These synergies lead to a higher yield of C2−C4 olefins in the joint transformation than theoretically expected by summing the yields of the individual reactions.16,17 Yan and Le Van Mao have also reported a higher yield of ethylene and propylene in the thermal-catalytic steam cracking of methanol/naphtha.26 Given the rapid deactivation of the catalyst, a further step toward the viability of the joint transformation of methanol and Received: Revised: Accepted: Published: 13073

May 2, 2012 July 13, 2012 September 13, 2012 September 13, 2012 dx.doi.org/10.1021/ie301142k | Ind. Eng. Chem. Res. 2012, 51, 13073−13084

Industrial & Engineering Chemistry Research

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particles are dried in an oven at 110 °C for 24 h and then calcined at 575 °C for 3 h. This temperature is reached following a ramp of 5 °C min−1. The agglomeration of the active phases does not significantly reduce acidity, although it improves the accessibility of the reactants (providing the catalyst with a matrix with mesopores and macropores), which is essential for reducing deactivation by coke deposition.18,27 Furthermore, agglomeration improves hydrothermal stability because zeolite crystals are more dispersed and the heat capacity of the material by mass unit of zeolite is higher.28 The catalyst allows performing up to 10 reaction− regeneration cycles, with no irreversible deactivation either in the reaction stage or in the regeneration stage, which is performed in situ by coke combustion with air in the reactor at 550 °C.29 Table 1 sets out the physical properties and the acidity values of the catalyst. The porous structure has been determined by

n-butane requires both a catalyst that can be regenerated and the knowledge of its reactivation kinetics, which will allow studying alternative reaction−regeneration systems. This paper addresses the reactivation of an HZSM-5 zeolite catalyst for which the process kinetics at zero time on stream and deactivation have already been determined.17 Furthermore, the operation in reaction−regeneration cycles has been studied in order to validate the simulation of these cycles based on process kinetic equations.

2. EXPERIMENTAL SECTION 2.1. Reaction Equipment and Product Analysis. The runs have been carried out under normal pressure in automated reaction equipment, as described in a previous paper.14 The reactor is made of 316 stainless steel with an internal diameter of 9 mm and 10 cm of effective length. It is located inside a ceramic covered stainless steel cylindrical chamber, which is heated by electric resistance and can operate at up to 100 atm and 700 °C with a catalyst mass of up to 5 g. The bed consists of a mixture of catalyst and inert solid (carborundum with an average particle diameter of 0.105 mm) to ensure bed isothermality and attain sufficient height under low space time conditions. The temperature is controlled by a digital TTM-125 series controller and measured by a thermocouple (K-type) situated in the catalyst bed. There are two other temperature controllers: one for the furnace chamber and the other for the transfer line between the reactor and the micro gas chromatograph (micro-GC). The operating variables are controlled by bespoke software. Bed isothermicity has been proven by longitudinally displacing the thermocouple. Isothermicity is enhanced by the fact that runs are carried out under energy-neutral process conditions. Reaction product samples (diluted in a He stream of 17 cm3 min−1) are continuously analyzed in a micro-GC (Varian CP4900). The remaining stream of reaction products passes through a Peltier cell at 0 °C. The amount of liquid condensate is controlled by a level sensor, and the noncondensable gas flow is vented. The micro-GC (with Star Toolbar software) is provided with three analytical modules and the following columns: Porapak Q (PPQ) (10 m), where the lighter products are separated (CO2, methane, ethane, ethylene, propane, propylene, methanol, dimethyl ether, water, butanes, and butenes); a molecular sieve (MS-5) (10 m), where H2, CO, O2, and N2 are separated; and 5CB (CPSIL) (8 m), where the C5−C10 fraction is separated. The compounds were identified and quantified based on calibration standards of known concentrations. The balance of atoms (C, H, O) is closed in all runs above 99.5%. 2.2. Catalyst. The catalyst has been selected in a previous paper, based on the combination of different criteria (activity at zero time on stream, olefin selectivity, deactivation by coke deposition, and hydrothermal stability).15 The selected catalyst (HZ-30) has been prepared using an HZSM-5 zeolite, with SiO2/Al2O3 = 30, supplied by Zeolyst International in ammonium form, which has been calcined at 570 °C in order to obtain the acid form. The zeolite has been agglomerated with a binder (bentonite, Exaloid) (30 wt %) and alumina (Prolabo, calcined at 1000 °C to render it inert) as inert charge (45 wt %). The catalyst particles have been obtained by wet extrusion through 0.8 mm diameter holes, using a high-pressure hydraulic piston. The extrudates are first dried at room temperature for 24 h and then sieved to a particle diameter between 0.15 and 0.3 mm. The

Table 1. Physical and Chemical Properties of the HZSM-5 Zeolite Catalyst acid strength, kJ (mol of NH3)−1 total acidity, (mmol of NH3)·g−1 dp, Å SBET, m2 g−1 Vm, cm3 g−1 Vp (17 < dp (Å) < 3000), cm3 g−1 pore volume distribution (%) dp (Å) < 20/20 < dp (Å) < 500/dp (Å) > 500 Brønsted/Lewis site ratio at 150 °C

120 0.23 102 220 0.044 0.69 2.96/46.5/50.5 1.50

N2 adsorption−desorption (Micromeritics ASAP 2010) and Hg porosimetry (Micromeritics Autopore 9220). The micropore volume corresponds to the active phase (HZSM-5 zeolite), whereas the volume of meso- and macropores corresponds to the catalyst matrix (bentonite and alumina). The total acidity and acid strength of the catalysts have been determined by monitoring the adsorption−desorption of NH3, by combining the techniques of thermogravimetric analysis and differential scanning calorimetry and temperature-programmed desorption using a Setaram TG−DSC calorimeter connected online with a Thermostar mass spectrometer from Balzers Instruments.30,31 The Brønsted/Lewis (B/L) acid site ratio (Table 1) has been determined by analyzing the region of 1400−1700 cm−1 in the Fourier transform infrared (FTIR) spectrum of adsorbed pyridine, which has been obtained using a Specac catalytic chamber connected online with a Nicolet 6700 FTIR spectrometer. The Brønsted/Lewis site ratio value at 150 °C has been determined from the ratio between the intensities of pyridine adsorption bands at 1545 and 1450 cm−1 and taking into account the molar extinction coefficients of both adsorption bands (εB = 1.67 cm μmol−1 and εL = 2.22 cm μmol−1).32

3. RESULTS 3.1. Operation in Reaction−Regeneration Cycles. The methodology for the kinetic study of catalyst reactivation by combusting the coke deposited in situ in the reactor has been detailed in a previous paper in which the transformation of methanol into hydrocarbons has been studied on an HZSM-5 zeolite catalyst by operating in reaction−regeneration cycles.33 The methodology consists of the following steps: (1) experimentation in reaction−regeneration cycles in a range of 13074

dx.doi.org/10.1021/ie301142k | Ind. Eng. Chem. Res. 2012, 51, 13073−13084


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Figure 1. Effect of coke combustion time on the evolution of reaction product yields and conversion in the second step with time on stream. Conditions in the first and second reaction steps: 500 °C; 0.57 gcatalyst h/molCH2. Coke combustion time: (a) 0, (b) 20, (c) 80, and (d) 120 min.

thereby making its combustion kinetics and combustion heat reproducible.34,35 The reaction products have been grouped into the following lumps: C2−C4 olefins (ethylene, propylene, and butenes); C2− C4 paraffins (ethane, propane, and isobutane); C5+ aliphatics (olefins and paraffins); aromatics (benzene, toluene, and xylenes) and methane. Regeneration Step. In the regeneration step, the conditions are the following: pressure, atmospheric; air flow rate, 30 mL min−1; combustion time at 550 °C, 0, 20, 40, 80, and 120 min. Figure 1 shows the results for the evolution of product lump yields and the conversion of the reactant mixture with time on stream. These results correspond to a reaction temperature of 500 °C. Each graph corresponds to a value of combustion time and shows the results for the first reaction step (with the fresh catalyst) and for the second step (with the catalyst regenerated for the corresponding combustion time). The yield of each lump has been determined as follows:

operating conditions (temperature and space time) for different lengths of time in the reaction and regeneration steps (time on stream and combustion time, respectively); (2) calculation of the kinetic parameters for the reactivation equation by fitting the experimental results for the evolution of component concentrations with time on stream (obtained with the catalyst under different partial reactivation states) to the values calculated by solving the component mass balances in the reactor. The kinetic models corresponding to zero time on stream and deactivation are required for solving these balances. The experimentation in reaction−regeneration cycles has been carried out under the following conditions. Reaction Step. For the reaction step, the conditions are as follows: pressure, atmospheric; methanol/n-butane ratio in the feed (CH2 equivalent units), 0.75 (corresponding to the regime of energy compensation between the individual reactions); temperature, 500 and 550 °C; space time, up to 0.57 gcatalyst h/ molCH2; length of the first and second steps (time on stream, t) up to 4 h. Subsequent to each reaction and prior to coke combustion, the deactivated catalyst bed has been swept with He (30 cm3 min−1) for 20 min at the reaction temperature. This sweeping treatment is aimed at (i) removing the adsorbed volatiles that do not deactivate the catalyst but affect its regeneration due to uncontrolled combustion and (ii) homogenizing the coke H/C ratio at all the longitudinal positions in the reactor and for any value of space time or time on stream.29 Consequently, the coke combustion kinetics is valid for all the positions in the reactor with differing coke content. It has been determined that a severe sweeping of the coke reduces its H/C ratio to values in the 0.5−0.7 range,

Yi =

Fi ·100 F0

(1)

where Fi and F0 are the molar flow rates of i lump at the outlet of the reactor and methanol + n-butane in the feed, respectively, expressed in CH2 equivalent units. The conversion of the methanol + n-butane combined feed has been defined as X= 13075

FS − F0 F0

(2)



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where FS is the flow rate of methanol + n-butane at the reactor outlet, measured in CH2 equivalent units. As observed in Figure 1a, the catalyst hardly recovers any activity when only sweeping with He at 550 °C (without coke combustion). Parts b and c of Figure 1 show partial activity recovery for combustion times of 20 and 80 min, respectively. Figure 1d shows that coke combustion for 120 min at 550 °C leads to the full recovery of catalyst activity and, therefore, the results for the second step (regenerated catalyst) are similar to those for the first step (fresh catalyst). It should be noted that the precaution has been taken of preparing the catalyst by calcination at 575 °C. The calcination at a temperature higher (25 °C above) than the maximum for reaction or coke combustion in the regeneration (550 °C in both cases) leads to an equilibrium state in the fresh catalyst in which the strong acid sites (presumably on the outer surface of the zeolite crystals) have been dehydroxylated. These sites undergo irreversible deactivation in the reaction step due to the high water content formed by methanol dehydration in the medium,30,31,36 or in the regeneration step due to hot spots.37 In view of the results in Figure 1d, catalyst deactivation by coke is reversible, given that the catalyst fully recovers its kinetic performance under all the operating conditions (temperature, space time, and time on stream) subsequent to a coke combustion step with air at 550 °C for 120 min. 3.2. Kinetic Modeling and Deactivation Kinetics. In the kinetic scheme in Figure 2, the components have been grouped

Reactions 2−4 are the formation of C2−C4 olefins from oxygenates (methanol and dimethyl ether) and n-butane. The most accepted mechanism assumes that the transformation of methanol into olefins occurs through polymethylbenzenes and methylbenzene-derived cations.24,39,40 Reactions 5 and 6 show the formation of CH4 (C) by the decomposition of oxygenates and n-butane. Reactions 7−11. They are autocatalytic reactions in which C2−C4 olefins are both reactants and products. Reactions 7 and 8 consider the activation of the hydrocarbon pool mechanism by the olefins in the reaction medium.24,39,40 Reactions 9−11 correspond to the formation of olefins by the oligomerization− cracking reactions of olefins with n-butane, C2−C4 paraffins (P), and C5−C10 fraction (lump G, made up of aromatics and nonaromatics).41 Reactions 12−14. They are reactions transforming C2−C4 olefins into n-butane and paraffins, mainly by hydrogen transfer, and into C 5−C 10 by oligomerization, cyclization, and aromatization.42 The coke that deactivates the catalyst is not considered in Figure 2 because its content is not significant in the mass balance. Coke deposition causes deactivation, which has been quantified by the following expression:17 da = −[kd1(yM + yD ) + kd2(yO + yG )]a dt

(4)

Equation 4 accounts for the dependence of the concentrations of oxygenates (methanol and dimethyl ether), C2−C4 olefins, and C5−C10 components on coke deposition. The expression is consistent with the well-reported role as coke precursors of oxygenates,18,38 olefins, and aromatics in the reaction medium.43 Selective deactivation has also been considered by acting on the reactions in the kinetic scheme. Thus, three levels of deactivation have been considered in the reactions of the kinetic scheme in Figure 2 (where the kinetic constant subscripts correspond to the reaction numbers): (i) severe deactivation in reactions 2−4 and 7−14, which are affected by coke deposition on the acid sites; (ii) moderate deactivation of thermal origin in reactions 5 and 6, with CH4 being formed, and with a secondary role played by acid sites; (iii) insignificant deactivation in the dehydration of methanol into dimethyl ether (reaction 1), which occurs on very weak acid sites. Consequently, the activity for t time on stream in the severe deactivation reactions is ri a= (ri)t = 0 (5)

Figure 2. Kinetic scheme of the joint transformation of methanol and n-butane.17

into lumps,17 given that they are more convenient for reactor design. The scheme integrates the catalytic transformation of methanol13 and n-butane14 in the same temperature range. In order to establish these schemes, well-reported mechanisms in the literature have been considered for the transformation of methanol and paraffins into olefins on an HZSM-5 zeolite. The kinetic scheme in Figure 2 includes reactions 1−14. Reaction 1 is the dehydration of methanol (M) to dimethyl ether (D) and water (W), through a mechanism with methoxy ion as intermediate, which is assumed to be in equilibrium, with the constant being38

for reactions 2−4 and 7−14 in the kinetic scheme in Figure 2. The activity in the reactions of moderate deactivation is ri a′ = a m = (ri)t = 0 (6) for reactions 5 and 6 in the kinetic scheme in Figure 2. In order to simplify the kinetic model, activity a′ has been related to activity a by means of exponent m (m < 1), which allows using only the parameter a for both activities. For reaction 1, in which deactivation is insignificant, a = constant = 1. According to the kinetic scheme in Figure 2 and the aforementioned considerations on deactivation, the kinetic equations for the formation of product lumps for any value of time on stream are as follows:

⎡ 1 K = exp⎢ −9.76 + 3200 + 1.07 log T − (0.66 × 10−3) ⎣ T 1 ⎤ T + (0.49 × 10−7)T 2 + 6500 2 ⎥ (3) T ⎦ 13076



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⎛k ⎞ rM = −k1yM 2 + ⎜ 1 yD yW ⎟ − (k 2yM + k 7yM yO )a − k5yM am ⎝K ⎠

Table 2. Kinetic Parameters for the Joint Transformation of Methanol and n-Butane17

(7)

rD = k1yM 2

⎛k ⎞ − ⎜ 1 yD yW ⎟ − (k 3yD + k 8yD yO )a − k5yD am ⎝K ⎠

kinetic constants at 500 °C k1*, molCH2 gcatal−1 h−1 (molM/mol)−2

(8)

rB = ( −k4yB − k 9yB yO + k12yO2 )a − k6yB am

k2*, molCH2 gcatal

(9)

rO = [k 2yM + k 3yD + k4yB + k 7yM yO + k 8yD yO + k 9yB yO + k10yP yO + k11yG yO − (k12 + k13 + k14)yO2 ]a 2

rP = ( −k10yP yO + k13yO )a 2

activation energies, kJ mol−1

(10)

rG = ( −k11yG yO + k14yO )a

(12)

rC = (k5yM + k5yD + k6yB )am

(13)

−5

E1

12.9

1.1 × 10

E2

4.04

1.6 × 10−5

E3

175

k4*, molCH2 gcatal−1 h−1 (molB/mol)−1

0.888

E4

0.312

k5*, molCH2 gcatal−1 h−1 (molC/mol)−1

0.101

E5

113

h

(molM/mol)

−1

−1

−1

0.011

E6

71.8

k7*, molCH2 gcatal−1 h−1 (molM molO/mol2)−1

18.9

E7

2.92

k8*, molCH2 gcatal−1 h−1 (molD molO/mol2)−1

187

E8

1.65

k9*,

The effect of the methanol/n-butane ratio in the feed on the coke content and nature has been studied in a previous paper.18 At 550 °C and with a space time of 2.4 gcatalyst h (molCH3)−1, the coke content in the catalyst for 5 h time on stream is 0.65 wt % (by mass unit of catalyst free of coke) for pure n-butane in the feed and up to 5.7 wt % for pure methanol in the feed. For the methanol/n-butane ratio of 3 used in this paper, the coke content is 4.5 wt %. These coke contents explain the significant catalyst deactivation. Nevertheless, coke yield (or feed percentage transformed into coke) is lower than 0.2 wt %. These coke yields are very low (much lower than those for the lumps considered) and have therefore been ignored in the kinetic equations, eqs 7−13. Furthermore, based on the results for coke deposition,18 selective deactivation is attributed to (i) the preferential deposition of coke on strong acid sites (required for reactions 2−4 and 7−14) and (ii) the fact that reactions 5 and 6 occur on weak sites. The corresponding kinetic parameters (values of the kinetic constants at 500 °C and activation energies) are set out in Table 2 (with m = 0.12). These results have been reported in a previous paper on the kinetic study of this process,17 in which a detailed explanation is given on the methodology for kinetic data analysis, model discrimination, and the calculation of kinetic parameters. The parameter m is empirical, and its low value indicates that the decrease in a′ with time on stream is much slower than the decrease in a. 3.3. Reactivation Kinetics. The reactivation kinetics has been determined using a routine developed in MATLAB, which solves the mass conservation equations for the components in the reaction medium. The code simultaneously solves the following: (i) the kinetics at zero time on stream; (ii) the equation for deactivation (in the reaction step); (iii) the reactivation kinetics (in the regeneration step). Ideal gas flow is assumed in the calculation. The calculation program requires the values of the independent variables corresponding to each experiment (temperature, space time, reaction step time, and regeneration step time), and there are several calculation steps for each set of experimental conditions (Figure 3). Step 1 is the simulation of the first reaction step (determination of the activity profile subsequent to the reaction step): the longitudinal profile of the activity remaining along

72.4

−1

k3*, molCH2 gcatal−1 h−1 (molD/mol)−1

k6*, molCH2 gcatal

(11)

−1

−1

molCH2 gcatal−1

h

−1

(molB/mol)

2 −1

0.109

E9

55.0

k10*, molCH2 gcatal−1 h−1 (molP molO/mol2)−1

1.37

E10

43.8

k11*, molCH2 gcatal−1 h−1 (molG molO/ mol2)−1 k12*, molCH2 gcatal−1 h−1 (molO/mol)−2

2603

E11

7.82

13.8

E12

77.2

k13*,

h

molCH2 gcatal−1

(molB molO/mol )

−1

−2

36.4

E13

4.15

k14*, molCH2 gcatal−1 h−1 (molO/mol)−2

718

E14

16.4

kd1*, h−1 kd2*, h−1 m OF

2.97 × 10−2 7.0 × 10−3 0.12 17.3

Ed1 Ed2

33.5 47.4

h (molO/mol)

the reactor, a, is calculated by solving the mass conservation equations for the components in the reaction using the kinetic model (eqs 7−13 and Table 2) with the initial condition that the activity is unity (fresh catalyst) at any longitudinal position. Step 2 is the simulation of the regeneration: the longitudinal profile of the activity subsequent to partial reactivation, a0, is calculated using a proposed reactivation kinetic equation. The reactivation equation has been established based on the coke combustion time (tc) according to the following expression: da 0 = k(1 − a0), dtc

where a0 = a for tc = 0

(14)

which after integration yields a0 = 1 − (1 − a) exp[−ktc]

(15)

The activity recovery rate during the regeneration step is a function of the residual activity level (at the end of the reaction or beginning of regeneration) at each longitudinal position in the bed. The parameter k quantifies this dependency, for which a linear relationship has been proven to be valid: k (min−1) = b1 + b2a

(16)

These expressions are empirical and have no direct relationship with deactivation kinetics, eq 4, but have proved to be valid for the description of an HZSM-5 zeolite catalyst reactivation in the transformation of methanol into hydrocarbons by operating in reaction−regeneration cycles.33 Step 3 is the simulation of the second reaction step: the evolution of product composition at the reactor outlet with time is calculated by solving the mass conservation equations for the components in the reaction using the kinetic model (Figure 2 and Table 2), with the initial condition that the activity at the different longitudinal positions in the reactor is 13077



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under the set of experimental conditions j, which is expressed as a molar fraction by mass unit of organic components in the reaction medium; Xi,j is the corresponding value calculated for i lump (following steps 1, 2, and 3 described above); nl is the number of lumps in the kinetic scheme; and p is the total number of experimental values. The weight factors for each lump have been calculated by the expression

ωi =

calculated with the reactivation equation (partially regenerated catalyst). Step 4. The program then minimizes an error objective function (OF) established as the weighted residual sum of squares, i.e., the sum of squared differences between the experimental and calculated values for product composition at the reactor outlet, corresponding to the second reaction step: nl

p

i=1

j=1

(18)

where nexp is the number of experimental runs, including repetitions. Thus, the minority lumps in the reaction medium have a higher weight fraction. Gayubo et al.44 have proved that the results provided by eq 18 are similar to those obtained from the distribution variances of the experimental results. It should be noted that the calculation program solves the mass balances in the reaction step by rigorously considering the past history of the catalyst, consistent with previous papers on the kinetic modeling of the deactivation of different catalysts in several reactions.38,45−47 Following the calculation procedure described, the optimum values have been determined for parameters b1 and b2 in eq 16: b1 = 4.22 × 10−2 and b2 = 1.69 × 10−2. Figures 4−6 compare the experimental results (points) with those calculated (lines) for the evolution of concentration of component lumps in the reaction medium with time on stream in the second reaction step. Each figure corresponds to a different catalyst deactivation severity. Thus, the results in Figure 4 correspond to reaction conditions under which deactivation is moderate (500 °C and 120 min time on stream), whereas Figures 5 and 6 correspond to cycles causing a more severe deactivation in the first reaction step (500 °C and 240 min time on stream (Figure 5) and 550 °C and 180 min time on stream (Figure 6). Each graph corresponds to a different combustion time (20 and 80 min in parts a and b, respectively, in Figures 4−6). Figure 7 shows the decrease in activity with time on stream calculated for the first reaction step at three longitudinal positions in the reactor, corresponding to the inlet, halfway between the inlet and the outlet, and the outlet. Figure 7a corresponds to the reaction at 500 °C, and Figure 7b corresponds to the reaction at 550 °C. Figure 8 shows activity recovery with combustion time at three longitudinal positions in the reactor. Figure 8a corresponds to the reactivation subsequent to the reaction step at 500 °C for a time on stream of 120 min. In Figure 8b, the reaction step has been carried out at 550 °C for a time on stream of 180 min, and consequently, the residual activity of the catalyst at the beginning of the reactivation is lower than in Figure 8a. The results calculated for product composition at the reactor outlet for the second reaction cycle are shown in Figures 4 and 6, and they correspond to the catalysts whose profile for the activity recovered in the regeneration is that shown in Figure 8a and 8b, respectively. Furthermore, Figures 7 and 8 evidence the significance of considering the longitudinal profile of activity in the reactor for the deactivation and reactivation steps, given that these profiles are very pronounced due to the considerable effect of the composition in the reaction medium on catalyst deactivation (eq 4). As observed in Figure 7, deactivation is faster at the inlet of the catalytic bed and the deactivation rate attenuates as the

Figure 3. Steps 1−3 for calculating the catalyst reactivation kinetics by experimentation in reaction−regeneration cycles.

OF =

1 n ∑ j =exp1 Xi

∑ ωi ∑ (Xi*,j − Xi ,j)2 (17)

where ωi is the weight factor of each i lump in the kinetic scheme; Xi,j* is the experimental value of conversion for i lump 13078



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Figure 5. Comparison between the experimental values (points) and those calculated (lines) for the evolution of product concentration in the second reaction step with time on stream. Reaction conditions: 500 °C and 0.57 gcatalyst h/molCH2; time on stream in the first reaction step, 240 min. Coke combustion time (at 550 °C with air): (a) 20 and (b) 80 min.


longitudinal position in the reactor is higher, given that the kinetic model of deactivation for this process considers the main precursors for coke formation to be the oxygenates (MeOH + DME), whose concentration decreases as the longitudinal position in the reactor is higher.16 Furthermore, Figure 8 shows a rapid activity recovery in this process for relatively low values of combustion time, even for a severely deactivated catalyst (Figure 8b). It is observed that the catalyst activity is higher than 0.95 for a combustion time of 80 min and the reactivation rate significantly attenuates for higher times. This trend in reactivation kinetics, which is quantified by means of eqs 7 and 8, is characteristic of regeneration by coke combustion in mesoporous catalysts and is attributed to the heterogeneous combustion of the coke located at different positions in the micropores and to the difficult combustion of the coke remaining at positions of difficult access for the air stream in the crystalline structure.29,48 3.4. Optimization of Reaction−Regeneration Cycles. A program has been developed in MATLAB for simulating the operation under reaction−regeneration cycles. It handles the kinetic equations of the main reaction, deactivation, and reactivation, and its purpose is to determine the operating conditions that maximize the production of olefins in the joint transformation of methanol and n-butane. The methodology used for the simulation is similar to that described previously for calculating the reactivation kinetics and consists of the following steps.

(i) The first reaction step is simulated for a given set of experimental conditions (temperature, space time, time on stream) by solving the corresponding mass conservation equations of the lumps of components in the fixed bed reactor together with the deactivation equation and considering the initial conditions in the reactor (a(ξ,t=0) = 1, fresh catalyst). Following this procedure, both the evolution of product lump composition with time on stream at the reactor outlet and the activity profile along the longitudinal position of the catalytic bed at the end of the first reaction step are calculated. (ii) Next is the calculation of the activity recovered at each longitudinal position of the catalytic bed during the regeneration step, for a given value of combustion time, by means of the reactivation equations, eqs 15 and 16. (iii) The second reaction step is simulated by solving the same set of equations as in the first reaction step, but taking the activity profile calculated in step ii (recovered activity) as the initial condition for solving the deactivation equation. (iv) Steps ii and iii are repeated successively until steady state is reached (constant activity profile at the end of the nth reaction step). Using the values of olefin composition (expressed as molar fraction by mass unit of organic components, XO) at the outlet of the reactor in the nth reaction step, the average olefin production rate (function to be maximized) is calculated as follows: 13079



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rp,O ̅ =

X̅ Ot ·100 (W /F0)(t + tb + tc)

Figure 7. Evolution of activity with time on stream at the bed inlet, at the midpoint between the inlet and the outlet and at the outlet in the first reaction step for different experimental conditions: (a) 500 °C, 0.57 gcatalyst h/molCH2, and 120 min time on stream; (b) 550 °C, 0.57 gcatalyst h/molCH2, and 180 min time on stream.

depending on temperature and space time in the reaction step and, consequently, an optimization including these variables would be overly complex and has not been considered. Table 3 sets out the maximum values of the apparent olefin production rate (in reaction−regeneration cycles) for the different operating conditions in the reaction step (temperature and space time), together with the corresponding values of time on stream in the reaction step and coke combustion time required for this peak production. It also shows the values of the average olefin molar fraction obtained by operating in reaction−regeneration cycles. Given that activity values vary according to the position in the reactor, the values of the activity in Table 3 (a0 for ξ = 0.5) correspond to the activity recovered subsequent to the regeneration at a position halfway between the inlet and the outlet in the catalytic bed, for the corresponding optimum values of reaction time and regeneration time. As in Figures 9 and 10, the results in Table 3 show that, for any value of space time, the maximum average olefin production rate is obtained at 500 °C. For a given temperature, the average production rate peaks for a low value of space time. Furthermore, the optimum reaction time decreases as the temperature is increased and the space time is decreased. This trend in the results is explained by the considerable activation energy of the deactivation equation constants and, therefore, the increase in the deactivation rate with temperature, which

(19)

where X̅ O is the average value of the molar fraction of the olefin lumps in the reaction medium (by mass unit of organic components) at the outlet of the reactor for the corresponding reaction cycle. It is calculated by integrating the evolution of the molar fraction, XO, with time on stream. t is the time on stream for the reaction step; tb is the sweeping time of the deactivated catalyst prior to coke combustion, for which a constant value of 20 min has been considered. This is the minimum time for coke structure homogenization, which is required for reproducible combustion and stable regeneration temperature.34,35 tc is the combustion time in the regeneration step, and W/F0 is the space time in the reaction step. By operating in reaction−regeneration cycles under given conditions of temperature and space time in the reaction step, the apparent olefin production rate is maximized by establishing the optimum values of time on stream in the reaction step (t) and combustion time in the regeneration step (tc). As an example of the results of the simulation program, Figures 9 and 10 show the contour maps of the average olefin production rate for the reaction temperatures of 500 and 550 °C, respectively. The results correspond to different combinations of reaction and combustion times. Each graph corresponds to a given value of space time. As observed, the values of t and tc required for obtaining the maximum olefin production rate are different 13080



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Figure 8. Activity recovered at different positions in the catalytic bed as a function of coke combustion time and for different conditions in the first reaction step: (a) 500 °C, 0.57 gcatalyst h/molCH2, and 120 min time on stream; (b) 550 °C, 0.57 gcatalyst h/molCH2, and 180 min time on stream. Figure 9. Contour maps of the average olefin production rate (mol/ gcatalyst h) at 500 °C for different combinations of reaction and coke combustion times. (a) 0.37 and (b) 0.57 gcatalyst h/molCH2.

requires regeneration to be carried out subsequent to shorter times on stream. Furthermore, olefins are intermediate products in the reaction kinetic scheme (Figure 2), whose formation rate increases rapidly with space time but then decreases for higher values of this parameter due to secondary reactions.16 Consequently, the maximum average production rate (defined by mass unit of catalyst, eq 19) is obtained for a very low value of space time. Furthermore, oxygenates (MeOH/DME) are the main precursors in deactivation by coke and, consequently, the deactivation rate is higher as space time is lower (reaction in an earlier state), which is consistent with the result that the optimum reaction time is shorter as space time is smaller. It should be noted that the optimum operating conditions do not correspond to the complete recovery of catalyst activity. This result has also been obtained for the operation in reaction−regeneration cycles in the isomerization of cisbutene,49 and in the MTG process in a fixed bed and in a fluidized bed with catalyst circulation.50,51 The optimum activity of the partially generated catalyst, a0, is slightly lower when the temperature is higher and decreases when the space time is 0.17 gcatalyst h/molCH2.

reaction and deactivation by coke (detailed in previous papers) allows simulating the joint transformation of methanol and nbutane into olefins by operating in reaction−regeneration cycles with the reaction in the 500−550 °C range and the regeneration by combustion with air at 550 °C. The rigorous methodology used is especially suitable for reactions with fast deactivation by coke, whose kinetics depends on the concentration of reaction medium components. The simulation of the joint transformation in reaction− regeneration cycles has allowed determining an optimum temperature of 500 °C for maximizing the average production rate of C2−C4 olefins. An average olefin production rate of 22 mol (h gcatalyst)−1 is obtained at this temperature with a space time of 0.37 gcatalyst h mol−1. Furthermore, the values for the optimum conditions (time on stream in the reaction step and time of coke combustion with air) increase as the temperature is decreased and the space time in the reaction step is increased. The values of time on stream and time of coke combustion corresponding to the aforementioned optimum temperature and space time are 40 and 15 min, respectively. Although the kinetic tools have been used in the design of the process in a fixed bed reactor, they are also applicable to the simulation of other reaction−regeneration systems, such as

4. CONCLUSIONS The methodology developed for determining catalyst reactivation by coke combustion has proved to be valid for obtaining a kinetics that combined with those for the main 13081



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Notes

The authors declare no competing financial interest.

■

ACKNOWLEDGMENTS This work has been carried out with the financial support of the Department of Education Universities and Research of the Basque Government (Project GIC07/24-IT-220-07), of the University of the Basque Country (UFI 11/39), and of the Ministry of Science and Innovation of the Spanish Government (Project CTQ2007-66571/PPQ).

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Figure 10. Contour maps of the average olefin production rate (mol/ gcatalyst h) at 550 °C for different combinations of reaction and coke combustion times. (a) 0.57 and (b) 0.77 gcatalyst h/molCH2.

Table 3. Optimum Conditions for Maximizing the Average Olefin Production Rate in the Operation in Reaction− Regeneration Cycles, for Different Temperature and Space Time Values T, °C 500

550

W/F0, gcatalyst h/ molCH2

t, min

tc, min

rp̅ ,O, mol/h gcatalyst

X̅ O

a0(ξ=0.5)

0.07 0.17 0.37 0.57 0.77 0.07 0.17 0.37 0.57 0.77

20 25 40 75 110 20 25 30 50 70

14 14 15 18 18 15 15 16 16 18

7 16 22 20 18 5 11 14 14 12

0.013 0.073 0.153 0.172 0.186 0.010 0.051 0.114 0.137 0.143

0.650 0.670 0.580 0.470 0.420 0.570 0.600 0.560 0.450 0.370

interconnected fluidized bed units, which may operate with the values calculated for the reaction and regeneration times.

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AUTHOR INFORMATION

Corresponding Author

*Tel.: 34-94-6015361. E-mail: [email protected]. 13082

NOMENCLATURE a = activity for reactions 2−4 and 7−14 in the kinetic scheme in Figure 2 a′ = activity for reactions 5 and 6 in the kinetic scheme in Figure 2 a0 = activity subsequent to partial reactivation, defined as the ratio of reaction rates b1, b2 = constants in the reactivation kinetic equation (eq 16) dp = pore diameter, Å Ej, Edj = activation energy for each kinetic constant in the kinetic scheme and for each deactivation kinetic constant, respectively, kJ mol−1 F0, FS = molar flow rates of oxygenates + n-butane at the inlet and outlet of the reactor, respectively, mol of CH2 equivalent units h−1 Fi = molar flow rate of i lump in the outlet stream, mol of CH2 equivalent units h−1 K = thermodynamic equilibrium constant in MeOH to DME dehydration (eq 3) k = constant in the reactivation kinetic equation (eq 14), min−1 kd = deactivation constant (eq 4), h−1 m = exponent (eq 6) n1 = number of lumps nexp = number of experimental runs, including repetitions OF = error objective function (eq 17) p = number of experimental runs ri = reaction rate equation for the formation of i lump in the kinetic scheme at t time on stream (reactions 7−13 in Figure 2), (moli)CH2 (gcatalyst h)−1 rp̅ ,O = average olefin production rate in a reaction− regeneration cycle (eq 19), (molO)CH2 (gcatalyst h)−1 SBET = BET surface area, m2 g−1 T = temperature, K t, tb, tc = time on stream, deactivated catalyst sweeping time, and coke combustion time, min Vm, Vp = micropore volume and pore volume, respectively, cm3 g−1 W = catalyst mass, g X = conversion of methanol/n-butane (eq 2) Xi = molar fraction of i lump by mass unit of organic compounds, in CH2 equivalent units Xi,j, X*i,j = calculated value of each i lump composition for the experimental set of conditions j, and average value determined with experiments repeated under the same set of experimental conditions j, in CH2 equivalent units XO = olefin molar fraction in a reaction step, in CH2 equivalent units X̅ O = average olefin molar fraction in a reaction step, in CH2 equivalent units dx.doi.org/10.1021/ie301142k | Ind. Eng. Chem. Res. 2012, 51, 13073−13084


Article

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yi = molar fraction of i lump, by mass unit of all the components in the reaction medium Yi = yield of i lump (eq 1) Abbreviations

B, C, D, G, O, P, M, W = n-butane, methane, dimethyl ether, C5−C10 lump, C2−C4 olefins, C2−C4 paraffins without nbutane, methanol, and water, respectively Greek Symbols

ξ = longitudinal position along the reactor, dimensionless ωi = weight factor for each i lump in eqs 17 and 18

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