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Al2O3 catalyst
Applied Catalysis A: General 349 (2008) 156–164
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Applied Catalysis A: General journal homepage: www.elsevier.com/locate/apcata
Transient kinetic modelling of propane dehydrogenation over a Pt–Sn–K/Al2O3 catalyst M.P. Lobera, C. Te´llez, J. Herguido, M. Mene´ndez * Aragon Institute for Engineering Research (I3A), University of Zaragoza, 50009 Zaragoza, Spain
A R T I C L E I N F O
A B S T R A C T
Article history: Received 3 April 2008 Received in revised form 30 June 2008 Accepted 21 July 2008 Available online 30 July 2008
A complete kinetic model of propane dehydrogenation to produce propene over a Pt–Sn–K/Al2O3 catalyst was obtained. This has been investigated over the temperature range of 460–540 8C at atmospheric pressure. A Langmuir–Hinshelwood mechanism provides the best fit for propane dehydrogenation, while a monolayer–multilayer mechanism is proposed for modelling the coke formation. In addition, the reaction rate of coke formation and its influence on catalyst deactivation and subsequent regeneration have been studied. Finally, a suitable mathematical model is developed for simulating the process behaviour in a two-zone fluidized bed reactor (TZFBR). ß 2008 Elsevier B.V. All rights reserved.
Keywords: Reaction mechanism Propane dehydrogenation Pt–Sn–K/Al2O3 catalyst Catalyst deactivation Coke formation rates Reactor simulation
1. Introduction The increasing demand for propylene derivatives such as polypropylene, acrylonitrile, propylene oxide, cumene, phenol, isopropylic alcohol and many others has produced a correspondingly heavy increase in propylene demand during the last 20 years. It is expected that in order to meet future propylene demand, other processes will be developed (dehydrogenation, oxidative dehydrogenation, metathesis or use of membrane reactors) in addition to traditional petrochemical or refinery processes (steam cracking, fluid catalytic cracking), since it is forecast that the propylene market will grow faster than the ethylene market [1]. Propane dehydrogenation is an equilibrium limited and highly 1 endothermic reaction (DHR298 K ¼ 120 kJ mol ) that is often carried out at high temperatures and atmospheric pressure using platinum or chromium catalysts. The reaction of olefins on platinum is faster than that of paraffins, given that olefins interact with platinum more strongly than paraffins. The role of platinum modifiers is to selectively weaken the platinum–olefin interaction without affecting the platinum–paraffin interaction. Arsenic, tin or germanium are among the metals reported as being platinum activity modifiers. The modifiers also improve stability against fouling by heavy carbonaceous materials. Platinum is a highly
* Corresponding author. Tel.: +34 976 761152; fax: +34 976 762142. E-mail address: [email protected] (M. Mene´ndez). 0926-860X/$ – see front matter ß 2008 Elsevier B.V. All rights reserved. doi:10.1016/j.apcata.2008.07.025
active catalytic element and is not required in large quantities for catalyzing the reaction when it is dispersed on a high surface area support. High dispersion is also necessary for achieving high selectivity to dehydrogenation relative to undesirable side reactions, such as cracking. The typical high surface area alumina supports employed have acidic sites that accelerate skeletal isomerisation, cracking, and polymerization of olefinic materials, and enhance coke formation. Furthermore, alkali or alkaline earth metals assist in the control of the acidity [2,3]. The formation of short chain hydrocarbons (cracking) and coke (coking) are the main undesired reactions associated with propane dehydrogenation. Furthermore, coke is rapidly formed at high temperatures and as a consequence the catalyst is deactivated and needs to be regenerated. These difficult process features require dedicated reactor technology. During the last decade, our group has been researching process alternatives. The objective is to employ a single vessel that could be continuously operated. To this end, we have developed different types of fluidized bed reactors that use separated oxygen and hydrocarbon feeds. The simplest version is called the two-zone fluidized bed reactor (TZFBR). It consists of a fluidized bed where oxygen is fed to the lower part of the reactor, mixed with an inert gas, and the hydrocarbon is fed at an intermediate point of the bed. In this way two zones are created in the reactor: in the lower zone the catalyst is reoxidized by gas phase oxygen, a process that causes the gas stream to become depleted in oxygen. In the upper zone, the chemical reaction takes place. Reaction products,
M.P. Lobera et al. / Applied Catalysis A: General 349 (2008) 156–164
Nomenclature a A CC Ci
activity reactor section (m2) coke concentration (mg coke (mg catalyst)1) concentration of a compound i in the bubble or emulsion phase (mol m3) monolayer coke (mg coke (mg catalyst)1) Cm multilayer coke (mg coke (mg catalyst)1) CM maximum coke concentration in mono layer Cmax (mg coke (mg catalyst)1) db bubble diameter (m) diffusion coefficient of compound i Di,gas Eai activation energy (kJ mol1) fraction of wake in bubbles fW Fi molar flow of a compound i (mol s1) g gravity constant (m2 s1) DHR reaction enthalpy (kJ mol1) reaction rate constant at T0 (mmol g1 min1bar1) k0i ki reaction rate constant (mmol g1 min1 bar1) kinetic constant of coke combustion (min1 bar1) kOX KB,E, KE,B gas exchange coefficient bubble–emulsion or emulsion–bubble (s1) KW,E, KE,W coefficient of solid exchange between wake– emulsion or emulsion–wake (s1) MSC model selection criteria number of carbon atoms of product i ni total flow rate (cm3(STP) min1) QT rate of coke formation (g s1 g1) RC rate of formation of a given compound i Ri (mol s1 kg1) propene selectivity SC3 H6 t time (min) T reaction temperature (K) reference temperature (K) T0 TZFBR two-zone fluidized bed reactor u0 gas velocity (cm3(STP) cm2 s1) ub bubble velocity (m s1) minimum fluidization velocity (cm3(STP) cm2 s1) umf relative gas velocity (u/umf) ur solid velocity (m s1) us W mass of catalyst (g) [X] concentration of X coke conversion XC propane conversion X C3 H8 z bed height (m)
Greek letters a fraction of bed in bubbles b heating rate (K min1) emf minimum fluidisation porosity gi parameters of deactivation model rcat catalysts density (kg m3) Subscripts E B W
in emulsion in bubble in wake
157
unconverted reactant and inert diluent leave together at the top of the bed. The good circulation of the solid, characteristic of fluidized beds, provides transport of solid between both zones and thus a stationary state is reached. In reactions where the problem is the fast deactivation of the catalyst by coking, the TZFBR may be employed to implement continuous catalyst regeneration in the reactor [4]. In a previous work, our group studied the dehydrogenation of propane over a commercial Cr2O3/Al2O3 catalyst [5]. This study focuses on improving the propene yield reached and on the development of new models of catalyst deactivation. Furthermore, we present a complete mechanistic kinetic model for the dehydrogenation, the coke formation and its influence on catalyst deactivation, and catalyst regeneration using a Pt–Sn–K/Al2O3 catalyst, in order to perform this reaction in a two-zone fluidized bed reactor. 2. Experimental 2.1. Catalyst preparation Pt–Sn–K/Al2O3 catalyst was prepared by incipient wet impregnation in accordance with the procedure described in the literature [6] but with variations in composition. Commercial g-Al2O3 support (Puralox Sasol Germany GmbH) with particle size 160– 250 mm, BET surface area 153 m2 g1, and pore volume 0.45 ml g1 was used with distilled water as the solvent. The nominal composition of the catalyst was 0.05, 0.14 and 0.10 wt.% of Pt, Sn and K, respectively. The catalyst was prepared by sequential impregnation. The support was first impregnated with 6 wt.% of SnCl22H2O (Sigma– Aldrich, 98% purity) dissolved into HNO3 (Sigma–Aldrich, 1N). It was then dried in air at 100 8C during 12 h and calcined in a fluidized bed reactor as follows. The catalyst was heated (3 8C min1) to 550 8C in air and calcined for 2.5 h at this temperature in humid air. The wet air was then replaced with dry air and the catalyst was calcined for a further 2 h at 550 8C. The purpose of the wet calcination was to reduce the chlorine content of the catalyst [6]. Pt was subsequently impregnated using H2PtCl66H2O (Sigma–Aldrich, 8 wt.% solution in water), and the drying and calcination were repeated. Finally, the K was impregnated with potassium nitrate solution (Sigma–Aldrich, 99% purity), and the drying and calcination were again repeated. The catalyst was ground and sieved to a particle size of 160– 250 mm. Before the kinetic studies, the catalyst was reduced in situ in a flow of H2:He = 1:2 during 2 h at 550 8C; reacted in C3H8:Ar = 1:1 during 30 min at 500 8C and regenerated in O2:He = 1:20 during 1 h at 500 8C. This reduction-reaction–regeneration cycle was repeated six times to obtain an aged and stable catalyst [7,8] with BET surface area 138 m2 g1 and pore volume 0.45 ml g1. 2.2. Thermogravimetric systems Coke formation experiments were performed in a CI Electronics thermobalance, under several different conditions. The reactor feed contained propane diluted in nitrogen. All the streams were mass flow controlled (Brooks). The range of operating conditions was as follows: temperature, 460–540 8C; molar fraction of propane, 0.25– 0.70; catalyst mass, 200 mg; particle size of 160–250 mm; atmospheric pressure and total flow rates of 300 cm3(STP) min1. A series of previous experiments was conducted varying the sample mass, particle size and flow rates in order to assure that no bed depth effect or diffusional limitations were present during the kinetic tests.
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Coke combustion experiments were performed in a MettlerToledo TGA/SDTA 851e micro-thermobalance, in non-isothermal conditions under a flow of air (50 cm3(STP) min1). The temperature was programmed from 30 to 600 8C with a heating rate of 5 8C min1. 2.3. Fluidized bed reactor The kinetic reaction study was carried out using a fluidized bed reactor inside an electrical furnace. The temperature was measured by a thermocouple inside a quartz thermowell, at the centre of the catalyst bed. A PID controller maintained temperature variations within 0.5 8C of the set point. The reactor was a 30-mmi.d., 300-mm-long quartz tube equipped with a quartz distributor plate. The reactor feed contained propane diluted in argon. All streams were mass flow controlled (Brooks). The range of operating conditions was as follows: temperature, 500–540 8C; percentage of propane in the reactor feed, 50–100%; catalyst mass, 10–30 g; relative gas velocity (ur = u0/umf), 1.2–4; a sieve fraction of 160–250 mm. The minimum fluidization velocity (umf) measured with He at 500 8C was 0.96 cm3(STP) cm2 s1. The experiments with the two-zone fluidized bed reactor were carried out in the installation that is described above. And TZFBR was 30-mm-i.d., 300-mm-long quartz tube equipped with a quartz distributor plate (for the Ar/O2 mixture). A movable axial quartz probe was used to introduce the propane at different reactor heights (ha). The reaction products were analyzed with an online gas chromatograph (Agilent 3000 mGC). Besides propene, cracked products are formed, mainly ethane, ethylene and methane. Propane conversion and propene selectivity were defined as follows: P X C3 H 8 ¼
Si ¼ P
i ðni =3Þ½F i out
P
½F C3 H8 out
i ðni =3Þ½F i out
100
ðni =3Þ½F i out 100 ½F C3 H8 out
i ðni =3Þ½F i out
(1)
(2)
where i includes all the components with carbon atoms in the exit gas stream, ni is the number of carbon atoms of component i and Fi is its molar flow. 2.4. Kinetic modelling The kinetic modelling has been developed following a procedure previously described [5]. Firstly, the data obtained in the fluidized bed reactor were fitted to a power-law model and different mechanistic models assuming a simple deactivation model. Secondly, thermogravimetric experiments were fitted to different coke formation models. Finally, the best kinetic models obtained previously were used to calculate the effect of coke formation on catalyst deactivation. Statistical parameters of the fitting were used to discriminate between the proposed models in all the steps. The fluidized bed reactor was assumed to be a plug flow reactor and solved by a finite difference approach. An isothermal bed was assumed due to the high-solid mixing. Fittings were performed with Scientist1 software that uses the Levenberg–Mardquardt algorithm. To obtain the optimal values of the rate parameters in the kinetic model, by comparing the experimental data with the model predictions, the least squares method is used. The kinetic parameters have been determined with a 95% confidence level. Once the kinetic has been calculated, the suitability of the models for representing the kinetic data can be assessed. To choose
the best kinetics model among those proposed in all the steps, the model selection criteria (MSC) [9] was used. It is useful because it takes into account the number of parameters of a given model and, therefore, allows a comparison of different models. The MSC has been normalized to be independent of the scale of data. Thus, the most appropriate model is one that has the highest value of MSC. And it is computed as: 2P
MSC ¼
l j¼1 ðY obs j ln4Pl j¼1 ðY obs j
2
Y¯obs Þ
Y cal j Þ2
3
5 2p l
(3)
where l is the number of experimental points, p the number of parameters and Y¯obs the weighted mean of the experimental observations. 3. Experimental results 3.1. Coke formation Coke formation was measured as a function of time with thermogravimetric experiments. Fig. 1 displays typical coke content versus time-on-stream at different temperatures. The coke content increases abruptly with time-on-stream during approximately the first 10 min followed by a more moderate linear increase during the rest of the reaction stage. The final coke content and coke formation rate in both the linear and the abrupt zones increase with the temperature. These results suggest that in the initial period the coke deposition itself causes the deactivation of the coking reaction, and after this period a residual coking activity remains constant in the second zone. Similar trends have ˜ a et al. [7,8], and van Sint been observed by Gasco´n et al. [5], Pen et al. [10]. In the studied range, the influence of the propane concentration in propane–argon mixtures was not appreciable. After 90 min on stream, the coke content was 1.8, 2.0 and 1.6 mg coke g1 of catalyst for, respectively, 25, 50 and 70% of propane in the feed, with the other conditions being similar (W = 200 mg; QT = 300 cm3(STP) min1; T = 540 8C). The apparent reaction order of the coking process with respect to the propane concentration was very low. This effect has also been observed by Gasco´n et al. [5], van Sint et al. [10], and Dumez and Froment [11]. In general, the coke formation includes the adsorption of the hydrocarbons on the catalysts surface which gives rise to the formation of coke precursors and later on to coke. It may be hypothesized that due to the strong hydrocarbon adsorption, the hydrocarbon surface coverage is close to unity in all the studied experimental conditions
Fig. 1. Coke concentration in catalyst as a function of time for different temperatures. W = 200 mg; QT = 300 cm3(STP) min1; propane/argon ratio = 0.5.
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and cracking are very endothermic reactions. However, the activity initially decreases faster at high temperatures, due again to a faster coke formation. Therefore, propane conversion after 2 h on stream, in the studied range, is higher with the lower temperature. Selectivity to propene (Fig. 2a) decreases with temperature because its effect over the cracking reaction is stronger than over dehydrogenation. The influence of the feed composition is shown in Fig. 2b. Propane conversion is almost independent of propane concentration, indicating a kinetic order close to 1 for propane pressure. The propylene selectivity slightly decreases when propane concentration is increased, and catalyst deactivation is similar in all experiments performed. The propane conversion increases with spatial time (Fig. 2c) while the selectivity to propene remains nearly constant. This behaviour indicates that cracking products are mainly formed from propane. Fig. 2. (a) Conversion of propane (solid symbols) and selectivity to propene (open symbols) as a function of time for different temperatures: (&) 500 8C; (~) 520 8C; ($) 540 8C. W/FA0 = 74 s kg mol1; ur = 1.7; propane/argon ratio = 0.5 (the lines are a visual help). (b) Conversion of propane (solid symbols) and selectivity to propene (open symbols) as a function of time for different feed composition: (&) 50% C3H8; (~) 75% C3H8; ($) 100% C3H8. W = 10 g; ur = 1.7; T = 500 8C (the lines are a visual help). (c) Conversion of propane (solid symbols) and selectivity to propene (open symbols) as a function of time for different spatial time, W/FA0: (&) 74 s kg mol1; (~) 147 s kg mol1; ($) 220 s kg mol1. ur = 1.7; T = 500 8C; propane/argon ratio = 0.5 (the lines are a visual help).
4. Kinetic modelling results The proposed kinetic scheme for the propane reactions over the prepared Pt–Sn–K/Al2O3 catalyst is a parallel network: dehydrogenation reaction: C3 H8 $ C3 H6 þ H2
(I)
cracking reaction: and therefore the coke formation is independent of the propane concentration in the gas phase.
C3 H8 ! Cracking products
(II)
coke formation: C3 H8 ! Coke
3.2. Dehydrogenation of propane Fig. 2a shows the effect of temperature on propane conversion and propene selectivity. In all the cases, a decline in conversion and propylene selectivity can be observed over time. When comparing the coke content with propane conversion profiles as a function of time-on-stream, one can observe that the decrease is similar in both parameters, indicating that coke accumulation on the catalysts is responsible for the conversion decrease. At short times, the propane conversion increases with temperature, as may be expected since both dehydrogenation
(III)
4.1. Dehydrogenation of propane and cracking reaction Several models (Table 1) have been tested in order to obtain the best fit for the dehydrogenation reaction. The first corresponds to a power-law equation. The others are based on the mechanistic Langmuir–Hinshelwood model. LHHW-1 supposes that adsorption equilibrium constants for propane and propene have the same order of magnitude. LHHW-2 assumes that propane adsorption is negligible. After combining reaction and adsorption equilibrium
Table 1 Kinetic equations for the proposed reaction mechanisms for propane dehydrogenation, coke formation and catalyst deactivation Model
Kinetic equation
Reaction kinetics
Power law
PC H PH 2 6 ðrC3 H8 Þ ¼ k1 P C3 H8 3K eq
LHHW-1
rC3 H8 ¼
k1 K C H ðPC H ðK C H =K C H ÞðPC H PH =K eq ÞÞ 2 3 8 3 8 3 6 3 8 3 6 1þðP C H =K C H ÞþðP C H =K C H Þ
LHHW-2
rC3 H8 ¼
k1 P C H ðPC H PH =K eq Þ 2 3 8 3 6 1þðP C H =K C H Þ
3 8
3 6
3 8
6
dC C dt
¼ k1C ðC max C m Þ þ k2C C m ;
Model C2 (h = 2, n = 0)
dC C dt
¼ k1C ðC max C m Þ2 þ k2C ;
Model C3 (h = 2, n = 1)
dC C dt
2
C C ¼ C max 2 C C ¼ Cmax
¼ k1C ðC max C m Þ þ k2C C m ;
CC ¼
a ¼ ð1 aC m Þ2
3 6
3 6
Coke formation kinetics Model C1 (h = 1, n = 1)
Deactivation model Model D1
3
Cm C m þC M
Model D2
a ¼ ð1 g 1 C m Þ þ g 2
Model D3
a ¼ ð1 g 1 C m Þ þ g 2 C m e½g 3 ðC M =C m Þ
h
h
k1C k2C k1C
i
k1C t 1þC max k1C t
h
i
1 eðk1C tÞ þ k2C C max t þ k2C t
k1C t 2 Cmax 1þC max k1C t
i
k
k2C ln½1 þ k1C C max t þ k2C C max t 1C
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and performing the active sites balance, the equations shown in Table 1 for each mechanism are obtained. The reaction rate constants are described by the Arrhenius equation. This can be expressed in a reparameterised form as follows: E 1 1 ki ¼ k0i exp ai (4) R T T0 where Eai is the activation energy, T0 the reference temperature (equal to 793.15 K, which was the average temperature in the range investigated), R the gas constant (8.314 J mol1 K1), and k0i the kinetic constant at T0. In all cases, the cracking reaction has been fitted to a power-law rate expression with first order relative to propane, with constant activity. To choose between the different kinetic models proposed, all the kinetic data were fitted using a simple and empirical deactivation model. This avoids potential errors resulting from extrapolating data to initial times to obtain initial rates. The model combines an exponential activity–coke relationship and an empirical function of coke content with the time-on-stream: Dumez and Froment [11]: a ¼ expðBC C Þ
(5)
Voorhies [12]: C C ¼ Atp
(6)
Thus, the activity–time relationship is given by a ¼ expðlt p Þ
(7)
where l is given by l = l0 exp[(Eal/R)((1/T) (1/T0))] and p a fixed value (p = 0.5) which does not depend on temperature. The use of this activity–time relationship allows us to employ all the conversion-selectivity data and not only from the extrapolation to 0 time. As may be seen in Table 2, the best fitting obtained (best MSC value) was for the Langmuir–Hinshelwood mechanism (LHHW-2) that accounts for inhibition by propene due to adsorption. This is the chosen model and it agrees with the equation proposed by Wan and Chu [13], for a ZnO/silicalite catalyst and Gasco´n et al. [5], for a Cr2O3/Al2O3 catalyst. This mechanistic model produced a better fit than the empirical potential model. 4.2. Coke formation kinetics The model used in this work to describe coke formation is a simple mechanistic model called the monolayer–multilayer coke growth model (MMCGM). This model was first proposed by Nam and Kittrell [14], generalized by Corella and Monzo´n [15], and used successfully by several authors [7,8,10,14–18]. In this model, the Table 2 Model selection criteria for kinetic studies (Table 1) Model
MSC
R2
Reaction kinetics Power law LHHW-1 LHHW-2
5.65 5.79 5.91
0.9973 0.9977 0.9980
Coke formation kinetics Model C1 Model C2 Model C3
2.24 2.90 2.62
0.9916 0.9957 0.9942
Deactivation model Model D1 Model D2 Model D3
5.86 5.58 5.96
0.9979 0.9972 0.9982
rate of coke deposition with time is given by the sum of coke formation on the surface of the catalyst (monolayer coke) and the rate of the simultaneous multilayer coke deposition: dC C dC m dC M ¼ þ dt dt dt
(8)
The formation of monolayer coke is proportional to the fraction of sites uncovered on the first layer: rCm ¼
dC m ¼ k1C ðC max C m Þh dt
(9)
where Cm is the coke concentration in monolayer, Cmax is the maximum coke concentration in monolayer, k1C is a kinetic coefficient and h is the kinetic order for monolayer coke formation. The formation of multilayer can start as soon as there is monolayer coke available and is proportional to the fraction of sites covered on the monolayer: rCM ¼
dC M n ¼ k2C Cm dt
(10)
where CM is the coke concentration in multilayer, k2C is the kinetic coefficient and n is the kinetic order for coke formation in multilayer. Both kinetic constants (k1C and k2C) are functions of the temperature and of the presence of hydrocarbons. The effect of partial pressure of hydrocarbons has been proved null in Section 3.1. Regarding temperature, Arrhenius-type parameters are used for both monolayer coke (i = 1) and multilayer coke (i = 2) formation: E 1 1 kiC ¼ k0iC exp aiC (11) R T T0 The values tested for the kinetic orders (h and n) are shown in Table 1 together with integrated equations for these values. Model C2 shows the best MSC value (Table 2) and gives an excellent fit of the coke versus time data at the different temperatures studied (model prediction curves in Fig. 1). This model has been chosen for the following steps of the fitting. The parameters obtained (Table 4) for coke formation show several interesting features. The fitting obtained for the kinetic order for monolayer coke formation (h = 2) would support the results found by other authors [7,15] that the coke formation step involves two sites and also satisfactorily describes the remarkable initial increase in the coke formation rates versus time. The value of the kinetic order for multilayer coke formation (n = 0) is consistent with the experimental results since this value predicts a constant activity of the multilayer coke. The activation energy for monolayer coking (38.4 kJ mol1) is lower than for multilayer coking (125.5 kJ mol1), which could be predicted for coke, which forms directly on the catalyst surface where metals can promote coke formation. Similar results were obtained by van Sint et al. [10], for Pt/Al2O3 catalyst and Gasco´n et al. [5], for Cr2O3/Al2O3 catalyst. It should also be noted that, in the fitting, the maximum coke concentration in monolayer is fixed for all the temperatures investigated. This value is 1 mg coke g1 of catalysts. 4.3. Deactivation model Several authors have proposed different expressions in order to correlate activity with coke content in catalysts [5,7,11,19–21]. Having chosen the best kinetic mechanism and the coke formation kinetics, the next step was to find the relationship between the coke content in the catalyst and the propane dehydrogenation activity.
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Table 3 Proposed kinetic scheme Propane dehydrogenation C3H8 $ C3H6 + H2 r C3 H8 ¼ a
k1 ðP C3 H8 ðPC3 H6 P H2 =K eq ÞÞ ; 1 þ ðPC3 H6 =K C3 H6 Þ
a ¼ ð1 g 1 C m Þ þ g 2 C m e½g 3 ðC M =C m Þ ; g 1 ¼ g 01 e½Eag 1 =Rðð1=TÞð1=T 0 Þ ;
Coke formation C3H6 $ 3CH0.5 + 2.25H2
k1 ¼ k01 e½Ea1 =Rðð1=TÞð1=T 0 Þ ; K C3 H6 ¼ K 0 e½DH=Rðð1=TÞð1=T 0 Þ h i dC C k1C t 2 ¼ k1C ðC max C m Þ2 þ k2C ; C m ¼ Cmax ; C M ¼ k2C t; k1C ¼ k01C e½Ea1C =Rðð1=TÞð1=T 0 Þ ; k2C ¼ k02C e½Ea2C =Rðð1=TÞð1=T 0 Þ 1þC max k t dt
Cracking reaction C3H8 $ C2H4 + CH4
ðr 2 Þ ¼ k2 P C3 H8 ;
Ethylene hydrogenation C2H4 + H2 $ C2H6
ðr 3 Þ ¼ k3 P C2 H4 PH2 ;
1C
k2 ¼ k02 e½Ea2 =Rðð1=TÞð1=T 0 Þ k3 ¼ k03 e½Ea3 =Rðð1=TÞð1=T 0 Þ
Coke deposited on the catalyst surface covers part of the active sites, resulting in the reduction of catalyst activity. According to the kinetic scheme, only monolayer coke would promote deactivation. This would imply that after all the monolayer coke was formed, the catalyst would have some remaining activity. However, the experimental activity–time results prove that activity decreases continuously with time, even after the monolayer has been fully formed. This suggests that multilayer coke also deactivates. The existence of catalytic activity even after the surface has been covered by coke can be explained by some catalytic activity of the coke itself, as has been found by several authors [5,19,22]. Three deactivation models have been tested, the empirical expressions shown in Table 1, where g1 varies with temperature according to the Arrhenius equation and g2 and g3 are constants independent of the temperature. The first deactivation model (D1) considers that catalyst activity depends on active sites in the catalyst surface. The other models (D2 and D3) associate the catalyst activity with the active sites in the catalyst surface and the remaining activity of the monolayer coke, Cm. The best MSC value is obtained for model D3 (Table 2), and it can be seen (Fig. 3) that the predicted values of the time evolution of molar flow for propane, propene and cracking products have a good correlation with the experimental ones. Table 4 lists the parameters for the chosen model (Table 3). The equilibrium constant value obtained for propene adsorption is very high. Furthermore, the activation energy of thermal cracking is greater than for the propane dehydrogenation (Table 4), in agreement with what is expected for a gas phase phenomenon versus a catalytic process as has been reported previously [5]. A
small standard deviation of kinetic parameters is obtained due to the large quantity of experimental data, which contributes to more precise results from a statistical point of view. Furthermore, the overall quality of the fits to the experimental data can be judged from the parity plots for molar flow of propane and propene at the reactor exit (Figs. 4 and 5), and an acceptable agreement of experimental and calculated data was found in the range of operating conditions used in this study. 4.4. Coke combustion kinetics The catalyst regeneration was carried out by coke combustion in non-isothermal conditions. Traditionally, analysis methods for non-isothermal data have been used in reactions of thermal decomposition (solid–gas). In order to model the weight loss of the solid in these reactions, the following expression is used:
dC C ¼ kOX ðTÞCCm PO2 dt
where kOX(T) is given by kOX ðTÞ ¼ k0OX eðEa =Rðð1=TÞð1=T 0 ÞÞÞ , CC is the coke concentration in the catalyst (mg coke (mg catalyst)1), m is the kinetic order for coke combustion, PO2 is the oxygen partial pressure and if a temperature ramp of constant heating rate b (K min1) is used, the time–temperature relationship is T = T0 + bt, where T0 is the initial temperature. Assuming m = 1, considering 0 coke conversion as XC = (C0Ct)/C0 and k0OX ¼ PO2 k0OX ; the Table 4 Kinetic parameter values for the kinetic models chosen Parameter
Fig. 3. Correlation between experimental data (symbols) and model predictions (lines) for the temporal evolution of molar flow of propane, propene, hydrogen and cracking at the reactor exit from deactivation model D3 (mmol min1). W/ FA0 = 74 s kg mol1; ur = 1.7; propane/argon ratio = 0.5; T = 520 8C.
(12)
Value
S.D.
Propane dehydrogenation 0.5242 mmol g1 min1 bar1 k01 Ea1 34.57 kJ mol1 DH 85.817 kJ mol1 K0 3.46 k02 0.00465 mmol g1 min1 bar1 Ea2 137.31 kJ mol1 k03 0.000236 mmol g1 min1 bar1 Ea3 154.54 kJ mol1 g01 948.92 (g cat) (g coke)1 Eag1 9.61 kJ mol1 g2 399 (g cat) (g coke)1 g3 40.07
0.0.1641 9.13 22.46 0.35 0.00243 37.82 8.01E5 15.09 67.08 1.94 43.21 14.39
Coke formation k01C Ea1C k02C Ea2C Cmax
234 (mg cat) (mg coke)1 min1 38.43 kJ mol1 1.45E6 (mg coke) (mg cat)1 min1 125.51 kJ mol1 1.04E3 (mg coke) (mg cat)1
6.76 1.38 4.98E8 1.85 5.20E6
Combustion model 0 k0OX Ea
0.0357 min1 4.27 kJ mol1
0.0017 0.087
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162
Fig. 4. Correlation between experimental data and model prediction for the molar flow of propane at the reactor exit (mmol min1) in the range of operating conditions used.
Fig. 6. Coke conversion as a function of time in non-isothermal conditions.
The parameter values were obtained by fitting the experimental data of coke conversion. An acceptable agreement of experimental and calculated data was found (Fig. 6). Table 4 lists the parameters obtained.
kinetics data were obtained independently, as described above. In order to take into account the ascending flow of solid, a three-phase model is proposed that includes the emulsion, bubble and wake phases where the gas existing inside the bubble, including the wake, is perfectly mixed [23]. It has been assumed that the gas rises in the emulsion at the minimum fluidisation velocity and that the remaining gas rises with the bubbles. The pressure, temperature, bubble size, fraction of bed in bubbles and fraction of the bubble volume occupied by the wake were considered to be constant. Under steady-state conditions, the mass conservation equations for a given gaseous species i are the following: In the bubble and wake phase:
5. Simulation of a two-zone fluidized bed reactor
dF i;B ¼ K B;E Aða þ a f W ÞðC i;B C i;E Þ þ Ri;B rcat Aa f W ð1 emf Þ (14) dz
integrated equation has the following expression: 1 X C ¼ ð1
Z 0 X C0 Þ exp k0OX
t
exp 0
Ea 1 1 dt T 0 þ bt T r R
(13)
As stated in Section 1, one of the requirements for the kinetic model developed in this work was that it should be able to predict reasonably well the performance of a two-zone fluidized bed reactor (Fig. 7). The model used for the reactor simulation contains no adjustable parameters and is in steady-state condition. In the two-zone fluidized bed reactor propane and oxygen are fed at different levels, providing separated zones for the reaction and catalyst regeneration and, in spite of the continuous coke formation, steady state is achieved because a dynamic equilibrium between coke formation in the reducing zone and coke burning in the oxidizing zone. The
Fig. 5. Correlation between experimental data and model prediction for the molar flow of propene at the reactor exit (mmol min1) in the range of operating conditions used.
In the emulsion phase: dF i;E ¼ K E;B Að1 a a f W ÞðC i;E C i;B Þ þ Ri;E rcat Að1 a dz a f W Þ f W ð1 emf Þ
Fig. 7. Scheme of the two-zone fluidized bed reactor (TZFBR).
(15)
M.P. Lobera et al. / Applied Catalysis A: General 349 (2008) 156–164
The rising velocity for bubbles in a bubbling bed is calculated with the gas velocity (u0): 0:5
uB ¼ ðu0 umf Þ þ 0:711ðgdb Þ
(16)
The gas exchange coefficient between bubble and emulsion (KB,E) is calculated as [24]: ! 0:25 D0:5 umf i;gas g þ 5:85 (17) K B;E ¼ 4:5 db d1:25 b The bubble diameter is calculated as the average values given by the Mori and Wen equation [25] with the initial and final bubble diameters. The fraction of bubbles in the bed and the gas exchange coefficient between emulsion and bubble (KE,B) are calculated to assure mass balance. Regarding the flow of solid, the main variable is the coke concentration, which will vary along the bed. In this case, the mass conservation equations are as follows: In wake phase: us;W
dC C;W ¼ K W;E ðC C;W C C;E Þ þ RC;W dz
(18)
In the emulsion phase: us;E
dC C;E ¼ K E;W ðC C;W C C;E Þ þ RC;E dz
(19)
The wake velocity is assumed to be the bubble velocity, and the emulsion velocity is calculated to assure the mass balance. The exchange velocity of solid between wakes and emulsion is calculated with an exchange coefficient, using the following equation proposed by Lim et al. [26]: K W;E ¼
0:15 db
(20)
The mass balances constitute a system of ordinary differential equations at steady state. This set of equations was solved by a Runge–Kutta–Merson method, fitting the concentration of the coke in the bottom of the reactor that would close the total mass balance for all the elementary components (carbon and oxygen). A comparison between theoretical and experimental results is shown in Fig. 8. In the TZFBR, the percentage of oxygen in the feed to the reactor is one of the most important variables [5]. Fig. 8a
163
shows its effect for a set of experiments performed with a constant input of propane (306 cm3(STP) min1) and varying argon and oxygen fluxes, keeping its addition constant (306 cm3(STP) min1). It may be seen that the model predicts satisfactorily the experimental trends, although the propane conversion is slightly higher and propene selectivity slightly lower. This difference could be attributed to the assumption made in the model, where oxygen that does not react with coke is used to burn propane, and experimentally the formation of water has been observed due to the reaction of oxygen with hydrogen. Therefore, CO2 formation is overestimated and a higher propane conversion and lower selectivity to propene are predicted. This effect increases with the oxygen content. Finally, Fig. 8b shows the effect of the temperature. The simulation predicts a slight decrease of the conversion and a mild decrease of the selectivity with the temperature. 6. Conclusions A complete kinetic model of propane dehydrogenation over a Pt–Sn–K/g-Al2O3 catalyst has been developed. Extensive experimental work has been performed to obtain the kinetic parameters involved in the reaction, in the coke formation, in the activity–coke content relationship and in the catalyst regeneration. A Langmuir–Hinshelwood mechanism with strong adsorption of propene has been chosen as it provides the best fit to describe the reaction kinetics. A monolayer–multilayer coke deposition model successfully describes the coking behaviour of Pt–Sn–K/Al2O3 catalyst during the dehydrogenation of propane. The relationship between coke content in the catalyst and catalyst activity has also been studied and determined, with the activity being related with both monolayer and multilayer coke. A power-law model has been proposed to describe the catalyst regeneration. The complete model explains reasonably well the qualitative and quantitative observations presented in this work. A three-phase reactor model combined with the proposed kinetic model describes satisfactorily the conversion and propene selectivity found experimentally in the catalytic dehydrogenation of propane in a two-zone fluidized bed reactor. Moreover, it explains their variations with the operating variables. Acknowledgement The authors thank DGI (Spain) for financial support in the project CTQ2004-01721/PPQ. References
Fig. 8. (a) Influence of the % of oxygen upon propane conversion and propene selectivity in TZFBR reactor. T = 500 8C; W = 60 g; Qpropane = 306 cm3(STP) min1; Qargon+oxygen = 306 cm3(STP) min1. (b) Influence of the temperature upon propane conversion and propene selectivity. W = 35 g; Qpropane = 437.5 cm3(STP) min1; Qargon = 350 cm3(STP) min1; Qoxygen = 87.5 cm3(STP) min1.
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Contents lists available at ScienceDirect
Applied Catalysis A: General journal homepage: www.elsevier.com/locate/apcata
Transient kinetic modelling of propane dehydrogenation over a Pt–Sn–K/Al2O3 catalyst M.P. Lobera, C. Te´llez, J. Herguido, M. Mene´ndez * Aragon Institute for Engineering Research (I3A), University of Zaragoza, 50009 Zaragoza, Spain
A R T I C L E I N F O
A B S T R A C T
Article history: Received 3 April 2008 Received in revised form 30 June 2008 Accepted 21 July 2008 Available online 30 July 2008
A complete kinetic model of propane dehydrogenation to produce propene over a Pt–Sn–K/Al2O3 catalyst was obtained. This has been investigated over the temperature range of 460–540 8C at atmospheric pressure. A Langmuir–Hinshelwood mechanism provides the best fit for propane dehydrogenation, while a monolayer–multilayer mechanism is proposed for modelling the coke formation. In addition, the reaction rate of coke formation and its influence on catalyst deactivation and subsequent regeneration have been studied. Finally, a suitable mathematical model is developed for simulating the process behaviour in a two-zone fluidized bed reactor (TZFBR). ß 2008 Elsevier B.V. All rights reserved.
Keywords: Reaction mechanism Propane dehydrogenation Pt–Sn–K/Al2O3 catalyst Catalyst deactivation Coke formation rates Reactor simulation
1. Introduction The increasing demand for propylene derivatives such as polypropylene, acrylonitrile, propylene oxide, cumene, phenol, isopropylic alcohol and many others has produced a correspondingly heavy increase in propylene demand during the last 20 years. It is expected that in order to meet future propylene demand, other processes will be developed (dehydrogenation, oxidative dehydrogenation, metathesis or use of membrane reactors) in addition to traditional petrochemical or refinery processes (steam cracking, fluid catalytic cracking), since it is forecast that the propylene market will grow faster than the ethylene market [1]. Propane dehydrogenation is an equilibrium limited and highly 1 endothermic reaction (DHR298 K ¼ 120 kJ mol ) that is often carried out at high temperatures and atmospheric pressure using platinum or chromium catalysts. The reaction of olefins on platinum is faster than that of paraffins, given that olefins interact with platinum more strongly than paraffins. The role of platinum modifiers is to selectively weaken the platinum–olefin interaction without affecting the platinum–paraffin interaction. Arsenic, tin or germanium are among the metals reported as being platinum activity modifiers. The modifiers also improve stability against fouling by heavy carbonaceous materials. Platinum is a highly
* Corresponding author. Tel.: +34 976 761152; fax: +34 976 762142. E-mail address: [email protected] (M. Mene´ndez). 0926-860X/$ – see front matter ß 2008 Elsevier B.V. All rights reserved. doi:10.1016/j.apcata.2008.07.025
active catalytic element and is not required in large quantities for catalyzing the reaction when it is dispersed on a high surface area support. High dispersion is also necessary for achieving high selectivity to dehydrogenation relative to undesirable side reactions, such as cracking. The typical high surface area alumina supports employed have acidic sites that accelerate skeletal isomerisation, cracking, and polymerization of olefinic materials, and enhance coke formation. Furthermore, alkali or alkaline earth metals assist in the control of the acidity [2,3]. The formation of short chain hydrocarbons (cracking) and coke (coking) are the main undesired reactions associated with propane dehydrogenation. Furthermore, coke is rapidly formed at high temperatures and as a consequence the catalyst is deactivated and needs to be regenerated. These difficult process features require dedicated reactor technology. During the last decade, our group has been researching process alternatives. The objective is to employ a single vessel that could be continuously operated. To this end, we have developed different types of fluidized bed reactors that use separated oxygen and hydrocarbon feeds. The simplest version is called the two-zone fluidized bed reactor (TZFBR). It consists of a fluidized bed where oxygen is fed to the lower part of the reactor, mixed with an inert gas, and the hydrocarbon is fed at an intermediate point of the bed. In this way two zones are created in the reactor: in the lower zone the catalyst is reoxidized by gas phase oxygen, a process that causes the gas stream to become depleted in oxygen. In the upper zone, the chemical reaction takes place. Reaction products,
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Nomenclature a A CC Ci
activity reactor section (m2) coke concentration (mg coke (mg catalyst)1) concentration of a compound i in the bubble or emulsion phase (mol m3) monolayer coke (mg coke (mg catalyst)1) Cm multilayer coke (mg coke (mg catalyst)1) CM maximum coke concentration in mono layer Cmax (mg coke (mg catalyst)1) db bubble diameter (m) diffusion coefficient of compound i Di,gas Eai activation energy (kJ mol1) fraction of wake in bubbles fW Fi molar flow of a compound i (mol s1) g gravity constant (m2 s1) DHR reaction enthalpy (kJ mol1) reaction rate constant at T0 (mmol g1 min1bar1) k0i ki reaction rate constant (mmol g1 min1 bar1) kinetic constant of coke combustion (min1 bar1) kOX KB,E, KE,B gas exchange coefficient bubble–emulsion or emulsion–bubble (s1) KW,E, KE,W coefficient of solid exchange between wake– emulsion or emulsion–wake (s1) MSC model selection criteria number of carbon atoms of product i ni total flow rate (cm3(STP) min1) QT rate of coke formation (g s1 g1) RC rate of formation of a given compound i Ri (mol s1 kg1) propene selectivity SC3 H6 t time (min) T reaction temperature (K) reference temperature (K) T0 TZFBR two-zone fluidized bed reactor u0 gas velocity (cm3(STP) cm2 s1) ub bubble velocity (m s1) minimum fluidization velocity (cm3(STP) cm2 s1) umf relative gas velocity (u/umf) ur solid velocity (m s1) us W mass of catalyst (g) [X] concentration of X coke conversion XC propane conversion X C3 H8 z bed height (m)
Greek letters a fraction of bed in bubbles b heating rate (K min1) emf minimum fluidisation porosity gi parameters of deactivation model rcat catalysts density (kg m3) Subscripts E B W
in emulsion in bubble in wake
157
unconverted reactant and inert diluent leave together at the top of the bed. The good circulation of the solid, characteristic of fluidized beds, provides transport of solid between both zones and thus a stationary state is reached. In reactions where the problem is the fast deactivation of the catalyst by coking, the TZFBR may be employed to implement continuous catalyst regeneration in the reactor [4]. In a previous work, our group studied the dehydrogenation of propane over a commercial Cr2O3/Al2O3 catalyst [5]. This study focuses on improving the propene yield reached and on the development of new models of catalyst deactivation. Furthermore, we present a complete mechanistic kinetic model for the dehydrogenation, the coke formation and its influence on catalyst deactivation, and catalyst regeneration using a Pt–Sn–K/Al2O3 catalyst, in order to perform this reaction in a two-zone fluidized bed reactor. 2. Experimental 2.1. Catalyst preparation Pt–Sn–K/Al2O3 catalyst was prepared by incipient wet impregnation in accordance with the procedure described in the literature [6] but with variations in composition. Commercial g-Al2O3 support (Puralox Sasol Germany GmbH) with particle size 160– 250 mm, BET surface area 153 m2 g1, and pore volume 0.45 ml g1 was used with distilled water as the solvent. The nominal composition of the catalyst was 0.05, 0.14 and 0.10 wt.% of Pt, Sn and K, respectively. The catalyst was prepared by sequential impregnation. The support was first impregnated with 6 wt.% of SnCl22H2O (Sigma– Aldrich, 98% purity) dissolved into HNO3 (Sigma–Aldrich, 1N). It was then dried in air at 100 8C during 12 h and calcined in a fluidized bed reactor as follows. The catalyst was heated (3 8C min1) to 550 8C in air and calcined for 2.5 h at this temperature in humid air. The wet air was then replaced with dry air and the catalyst was calcined for a further 2 h at 550 8C. The purpose of the wet calcination was to reduce the chlorine content of the catalyst [6]. Pt was subsequently impregnated using H2PtCl66H2O (Sigma–Aldrich, 8 wt.% solution in water), and the drying and calcination were repeated. Finally, the K was impregnated with potassium nitrate solution (Sigma–Aldrich, 99% purity), and the drying and calcination were again repeated. The catalyst was ground and sieved to a particle size of 160– 250 mm. Before the kinetic studies, the catalyst was reduced in situ in a flow of H2:He = 1:2 during 2 h at 550 8C; reacted in C3H8:Ar = 1:1 during 30 min at 500 8C and regenerated in O2:He = 1:20 during 1 h at 500 8C. This reduction-reaction–regeneration cycle was repeated six times to obtain an aged and stable catalyst [7,8] with BET surface area 138 m2 g1 and pore volume 0.45 ml g1. 2.2. Thermogravimetric systems Coke formation experiments were performed in a CI Electronics thermobalance, under several different conditions. The reactor feed contained propane diluted in nitrogen. All the streams were mass flow controlled (Brooks). The range of operating conditions was as follows: temperature, 460–540 8C; molar fraction of propane, 0.25– 0.70; catalyst mass, 200 mg; particle size of 160–250 mm; atmospheric pressure and total flow rates of 300 cm3(STP) min1. A series of previous experiments was conducted varying the sample mass, particle size and flow rates in order to assure that no bed depth effect or diffusional limitations were present during the kinetic tests.
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Coke combustion experiments were performed in a MettlerToledo TGA/SDTA 851e micro-thermobalance, in non-isothermal conditions under a flow of air (50 cm3(STP) min1). The temperature was programmed from 30 to 600 8C with a heating rate of 5 8C min1. 2.3. Fluidized bed reactor The kinetic reaction study was carried out using a fluidized bed reactor inside an electrical furnace. The temperature was measured by a thermocouple inside a quartz thermowell, at the centre of the catalyst bed. A PID controller maintained temperature variations within 0.5 8C of the set point. The reactor was a 30-mmi.d., 300-mm-long quartz tube equipped with a quartz distributor plate. The reactor feed contained propane diluted in argon. All streams were mass flow controlled (Brooks). The range of operating conditions was as follows: temperature, 500–540 8C; percentage of propane in the reactor feed, 50–100%; catalyst mass, 10–30 g; relative gas velocity (ur = u0/umf), 1.2–4; a sieve fraction of 160–250 mm. The minimum fluidization velocity (umf) measured with He at 500 8C was 0.96 cm3(STP) cm2 s1. The experiments with the two-zone fluidized bed reactor were carried out in the installation that is described above. And TZFBR was 30-mm-i.d., 300-mm-long quartz tube equipped with a quartz distributor plate (for the Ar/O2 mixture). A movable axial quartz probe was used to introduce the propane at different reactor heights (ha). The reaction products were analyzed with an online gas chromatograph (Agilent 3000 mGC). Besides propene, cracked products are formed, mainly ethane, ethylene and methane. Propane conversion and propene selectivity were defined as follows: P X C3 H 8 ¼
Si ¼ P
i ðni =3Þ½F i out
P
½F C3 H8 out
i ðni =3Þ½F i out
100
ðni =3Þ½F i out 100 ½F C3 H8 out
i ðni =3Þ½F i out
(1)
(2)
where i includes all the components with carbon atoms in the exit gas stream, ni is the number of carbon atoms of component i and Fi is its molar flow. 2.4. Kinetic modelling The kinetic modelling has been developed following a procedure previously described [5]. Firstly, the data obtained in the fluidized bed reactor were fitted to a power-law model and different mechanistic models assuming a simple deactivation model. Secondly, thermogravimetric experiments were fitted to different coke formation models. Finally, the best kinetic models obtained previously were used to calculate the effect of coke formation on catalyst deactivation. Statistical parameters of the fitting were used to discriminate between the proposed models in all the steps. The fluidized bed reactor was assumed to be a plug flow reactor and solved by a finite difference approach. An isothermal bed was assumed due to the high-solid mixing. Fittings were performed with Scientist1 software that uses the Levenberg–Mardquardt algorithm. To obtain the optimal values of the rate parameters in the kinetic model, by comparing the experimental data with the model predictions, the least squares method is used. The kinetic parameters have been determined with a 95% confidence level. Once the kinetic has been calculated, the suitability of the models for representing the kinetic data can be assessed. To choose
the best kinetics model among those proposed in all the steps, the model selection criteria (MSC) [9] was used. It is useful because it takes into account the number of parameters of a given model and, therefore, allows a comparison of different models. The MSC has been normalized to be independent of the scale of data. Thus, the most appropriate model is one that has the highest value of MSC. And it is computed as: 2P
MSC ¼
l j¼1 ðY obs j ln4Pl j¼1 ðY obs j
2
Y¯obs Þ
Y cal j Þ2
3
5 2p l
(3)
where l is the number of experimental points, p the number of parameters and Y¯obs the weighted mean of the experimental observations. 3. Experimental results 3.1. Coke formation Coke formation was measured as a function of time with thermogravimetric experiments. Fig. 1 displays typical coke content versus time-on-stream at different temperatures. The coke content increases abruptly with time-on-stream during approximately the first 10 min followed by a more moderate linear increase during the rest of the reaction stage. The final coke content and coke formation rate in both the linear and the abrupt zones increase with the temperature. These results suggest that in the initial period the coke deposition itself causes the deactivation of the coking reaction, and after this period a residual coking activity remains constant in the second zone. Similar trends have ˜ a et al. [7,8], and van Sint been observed by Gasco´n et al. [5], Pen et al. [10]. In the studied range, the influence of the propane concentration in propane–argon mixtures was not appreciable. After 90 min on stream, the coke content was 1.8, 2.0 and 1.6 mg coke g1 of catalyst for, respectively, 25, 50 and 70% of propane in the feed, with the other conditions being similar (W = 200 mg; QT = 300 cm3(STP) min1; T = 540 8C). The apparent reaction order of the coking process with respect to the propane concentration was very low. This effect has also been observed by Gasco´n et al. [5], van Sint et al. [10], and Dumez and Froment [11]. In general, the coke formation includes the adsorption of the hydrocarbons on the catalysts surface which gives rise to the formation of coke precursors and later on to coke. It may be hypothesized that due to the strong hydrocarbon adsorption, the hydrocarbon surface coverage is close to unity in all the studied experimental conditions
Fig. 1. Coke concentration in catalyst as a function of time for different temperatures. W = 200 mg; QT = 300 cm3(STP) min1; propane/argon ratio = 0.5.
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and cracking are very endothermic reactions. However, the activity initially decreases faster at high temperatures, due again to a faster coke formation. Therefore, propane conversion after 2 h on stream, in the studied range, is higher with the lower temperature. Selectivity to propene (Fig. 2a) decreases with temperature because its effect over the cracking reaction is stronger than over dehydrogenation. The influence of the feed composition is shown in Fig. 2b. Propane conversion is almost independent of propane concentration, indicating a kinetic order close to 1 for propane pressure. The propylene selectivity slightly decreases when propane concentration is increased, and catalyst deactivation is similar in all experiments performed. The propane conversion increases with spatial time (Fig. 2c) while the selectivity to propene remains nearly constant. This behaviour indicates that cracking products are mainly formed from propane. Fig. 2. (a) Conversion of propane (solid symbols) and selectivity to propene (open symbols) as a function of time for different temperatures: (&) 500 8C; (~) 520 8C; ($) 540 8C. W/FA0 = 74 s kg mol1; ur = 1.7; propane/argon ratio = 0.5 (the lines are a visual help). (b) Conversion of propane (solid symbols) and selectivity to propene (open symbols) as a function of time for different feed composition: (&) 50% C3H8; (~) 75% C3H8; ($) 100% C3H8. W = 10 g; ur = 1.7; T = 500 8C (the lines are a visual help). (c) Conversion of propane (solid symbols) and selectivity to propene (open symbols) as a function of time for different spatial time, W/FA0: (&) 74 s kg mol1; (~) 147 s kg mol1; ($) 220 s kg mol1. ur = 1.7; T = 500 8C; propane/argon ratio = 0.5 (the lines are a visual help).
4. Kinetic modelling results The proposed kinetic scheme for the propane reactions over the prepared Pt–Sn–K/Al2O3 catalyst is a parallel network: dehydrogenation reaction: C3 H8 $ C3 H6 þ H2
(I)
cracking reaction: and therefore the coke formation is independent of the propane concentration in the gas phase.
C3 H8 ! Cracking products
(II)
coke formation: C3 H8 ! Coke
3.2. Dehydrogenation of propane Fig. 2a shows the effect of temperature on propane conversion and propene selectivity. In all the cases, a decline in conversion and propylene selectivity can be observed over time. When comparing the coke content with propane conversion profiles as a function of time-on-stream, one can observe that the decrease is similar in both parameters, indicating that coke accumulation on the catalysts is responsible for the conversion decrease. At short times, the propane conversion increases with temperature, as may be expected since both dehydrogenation
(III)
4.1. Dehydrogenation of propane and cracking reaction Several models (Table 1) have been tested in order to obtain the best fit for the dehydrogenation reaction. The first corresponds to a power-law equation. The others are based on the mechanistic Langmuir–Hinshelwood model. LHHW-1 supposes that adsorption equilibrium constants for propane and propene have the same order of magnitude. LHHW-2 assumes that propane adsorption is negligible. After combining reaction and adsorption equilibrium
Table 1 Kinetic equations for the proposed reaction mechanisms for propane dehydrogenation, coke formation and catalyst deactivation Model
Kinetic equation
Reaction kinetics
Power law
PC H PH 2 6 ðrC3 H8 Þ ¼ k1 P C3 H8 3K eq
LHHW-1
rC3 H8 ¼
k1 K C H ðPC H ðK C H =K C H ÞðPC H PH =K eq ÞÞ 2 3 8 3 8 3 6 3 8 3 6 1þðP C H =K C H ÞþðP C H =K C H Þ
LHHW-2
rC3 H8 ¼
k1 P C H ðPC H PH =K eq Þ 2 3 8 3 6 1þðP C H =K C H Þ
3 8
3 6
3 8
6
dC C dt
¼ k1C ðC max C m Þ þ k2C C m ;
Model C2 (h = 2, n = 0)
dC C dt
¼ k1C ðC max C m Þ2 þ k2C ;
Model C3 (h = 2, n = 1)
dC C dt
2
C C ¼ C max 2 C C ¼ Cmax
¼ k1C ðC max C m Þ þ k2C C m ;
CC ¼
a ¼ ð1 aC m Þ2
3 6
3 6
Coke formation kinetics Model C1 (h = 1, n = 1)
Deactivation model Model D1
3
Cm C m þC M
Model D2
a ¼ ð1 g 1 C m Þ þ g 2
Model D3
a ¼ ð1 g 1 C m Þ þ g 2 C m e½g 3 ðC M =C m Þ
h
h
k1C k2C k1C
i
k1C t 1þC max k1C t
h
i
1 eðk1C tÞ þ k2C C max t þ k2C t
k1C t 2 Cmax 1þC max k1C t
i
k
k2C ln½1 þ k1C C max t þ k2C C max t 1C
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and performing the active sites balance, the equations shown in Table 1 for each mechanism are obtained. The reaction rate constants are described by the Arrhenius equation. This can be expressed in a reparameterised form as follows: E 1 1 ki ¼ k0i exp ai (4) R T T0 where Eai is the activation energy, T0 the reference temperature (equal to 793.15 K, which was the average temperature in the range investigated), R the gas constant (8.314 J mol1 K1), and k0i the kinetic constant at T0. In all cases, the cracking reaction has been fitted to a power-law rate expression with first order relative to propane, with constant activity. To choose between the different kinetic models proposed, all the kinetic data were fitted using a simple and empirical deactivation model. This avoids potential errors resulting from extrapolating data to initial times to obtain initial rates. The model combines an exponential activity–coke relationship and an empirical function of coke content with the time-on-stream: Dumez and Froment [11]: a ¼ expðBC C Þ
(5)
Voorhies [12]: C C ¼ Atp
(6)
Thus, the activity–time relationship is given by a ¼ expðlt p Þ
(7)
where l is given by l = l0 exp[(Eal/R)((1/T) (1/T0))] and p a fixed value (p = 0.5) which does not depend on temperature. The use of this activity–time relationship allows us to employ all the conversion-selectivity data and not only from the extrapolation to 0 time. As may be seen in Table 2, the best fitting obtained (best MSC value) was for the Langmuir–Hinshelwood mechanism (LHHW-2) that accounts for inhibition by propene due to adsorption. This is the chosen model and it agrees with the equation proposed by Wan and Chu [13], for a ZnO/silicalite catalyst and Gasco´n et al. [5], for a Cr2O3/Al2O3 catalyst. This mechanistic model produced a better fit than the empirical potential model. 4.2. Coke formation kinetics The model used in this work to describe coke formation is a simple mechanistic model called the monolayer–multilayer coke growth model (MMCGM). This model was first proposed by Nam and Kittrell [14], generalized by Corella and Monzo´n [15], and used successfully by several authors [7,8,10,14–18]. In this model, the Table 2 Model selection criteria for kinetic studies (Table 1) Model
MSC
R2
Reaction kinetics Power law LHHW-1 LHHW-2
5.65 5.79 5.91
0.9973 0.9977 0.9980
Coke formation kinetics Model C1 Model C2 Model C3
2.24 2.90 2.62
0.9916 0.9957 0.9942
Deactivation model Model D1 Model D2 Model D3
5.86 5.58 5.96
0.9979 0.9972 0.9982
rate of coke deposition with time is given by the sum of coke formation on the surface of the catalyst (monolayer coke) and the rate of the simultaneous multilayer coke deposition: dC C dC m dC M ¼ þ dt dt dt
(8)
The formation of monolayer coke is proportional to the fraction of sites uncovered on the first layer: rCm ¼
dC m ¼ k1C ðC max C m Þh dt
(9)
where Cm is the coke concentration in monolayer, Cmax is the maximum coke concentration in monolayer, k1C is a kinetic coefficient and h is the kinetic order for monolayer coke formation. The formation of multilayer can start as soon as there is monolayer coke available and is proportional to the fraction of sites covered on the monolayer: rCM ¼
dC M n ¼ k2C Cm dt
(10)
where CM is the coke concentration in multilayer, k2C is the kinetic coefficient and n is the kinetic order for coke formation in multilayer. Both kinetic constants (k1C and k2C) are functions of the temperature and of the presence of hydrocarbons. The effect of partial pressure of hydrocarbons has been proved null in Section 3.1. Regarding temperature, Arrhenius-type parameters are used for both monolayer coke (i = 1) and multilayer coke (i = 2) formation: E 1 1 kiC ¼ k0iC exp aiC (11) R T T0 The values tested for the kinetic orders (h and n) are shown in Table 1 together with integrated equations for these values. Model C2 shows the best MSC value (Table 2) and gives an excellent fit of the coke versus time data at the different temperatures studied (model prediction curves in Fig. 1). This model has been chosen for the following steps of the fitting. The parameters obtained (Table 4) for coke formation show several interesting features. The fitting obtained for the kinetic order for monolayer coke formation (h = 2) would support the results found by other authors [7,15] that the coke formation step involves two sites and also satisfactorily describes the remarkable initial increase in the coke formation rates versus time. The value of the kinetic order for multilayer coke formation (n = 0) is consistent with the experimental results since this value predicts a constant activity of the multilayer coke. The activation energy for monolayer coking (38.4 kJ mol1) is lower than for multilayer coking (125.5 kJ mol1), which could be predicted for coke, which forms directly on the catalyst surface where metals can promote coke formation. Similar results were obtained by van Sint et al. [10], for Pt/Al2O3 catalyst and Gasco´n et al. [5], for Cr2O3/Al2O3 catalyst. It should also be noted that, in the fitting, the maximum coke concentration in monolayer is fixed for all the temperatures investigated. This value is 1 mg coke g1 of catalysts. 4.3. Deactivation model Several authors have proposed different expressions in order to correlate activity with coke content in catalysts [5,7,11,19–21]. Having chosen the best kinetic mechanism and the coke formation kinetics, the next step was to find the relationship between the coke content in the catalyst and the propane dehydrogenation activity.
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Table 3 Proposed kinetic scheme Propane dehydrogenation C3H8 $ C3H6 + H2 r C3 H8 ¼ a
k1 ðP C3 H8 ðPC3 H6 P H2 =K eq ÞÞ ; 1 þ ðPC3 H6 =K C3 H6 Þ
a ¼ ð1 g 1 C m Þ þ g 2 C m e½g 3 ðC M =C m Þ ; g 1 ¼ g 01 e½Eag 1 =Rðð1=TÞð1=T 0 Þ ;
Coke formation C3H6 $ 3CH0.5 + 2.25H2
k1 ¼ k01 e½Ea1 =Rðð1=TÞð1=T 0 Þ ; K C3 H6 ¼ K 0 e½DH=Rðð1=TÞð1=T 0 Þ h i dC C k1C t 2 ¼ k1C ðC max C m Þ2 þ k2C ; C m ¼ Cmax ; C M ¼ k2C t; k1C ¼ k01C e½Ea1C =Rðð1=TÞð1=T 0 Þ ; k2C ¼ k02C e½Ea2C =Rðð1=TÞð1=T 0 Þ 1þC max k t dt
Cracking reaction C3H8 $ C2H4 + CH4
ðr 2 Þ ¼ k2 P C3 H8 ;
Ethylene hydrogenation C2H4 + H2 $ C2H6
ðr 3 Þ ¼ k3 P C2 H4 PH2 ;
1C
k2 ¼ k02 e½Ea2 =Rðð1=TÞð1=T 0 Þ k3 ¼ k03 e½Ea3 =Rðð1=TÞð1=T 0 Þ
Coke deposited on the catalyst surface covers part of the active sites, resulting in the reduction of catalyst activity. According to the kinetic scheme, only monolayer coke would promote deactivation. This would imply that after all the monolayer coke was formed, the catalyst would have some remaining activity. However, the experimental activity–time results prove that activity decreases continuously with time, even after the monolayer has been fully formed. This suggests that multilayer coke also deactivates. The existence of catalytic activity even after the surface has been covered by coke can be explained by some catalytic activity of the coke itself, as has been found by several authors [5,19,22]. Three deactivation models have been tested, the empirical expressions shown in Table 1, where g1 varies with temperature according to the Arrhenius equation and g2 and g3 are constants independent of the temperature. The first deactivation model (D1) considers that catalyst activity depends on active sites in the catalyst surface. The other models (D2 and D3) associate the catalyst activity with the active sites in the catalyst surface and the remaining activity of the monolayer coke, Cm. The best MSC value is obtained for model D3 (Table 2), and it can be seen (Fig. 3) that the predicted values of the time evolution of molar flow for propane, propene and cracking products have a good correlation with the experimental ones. Table 4 lists the parameters for the chosen model (Table 3). The equilibrium constant value obtained for propene adsorption is very high. Furthermore, the activation energy of thermal cracking is greater than for the propane dehydrogenation (Table 4), in agreement with what is expected for a gas phase phenomenon versus a catalytic process as has been reported previously [5]. A
small standard deviation of kinetic parameters is obtained due to the large quantity of experimental data, which contributes to more precise results from a statistical point of view. Furthermore, the overall quality of the fits to the experimental data can be judged from the parity plots for molar flow of propane and propene at the reactor exit (Figs. 4 and 5), and an acceptable agreement of experimental and calculated data was found in the range of operating conditions used in this study. 4.4. Coke combustion kinetics The catalyst regeneration was carried out by coke combustion in non-isothermal conditions. Traditionally, analysis methods for non-isothermal data have been used in reactions of thermal decomposition (solid–gas). In order to model the weight loss of the solid in these reactions, the following expression is used:
dC C ¼ kOX ðTÞCCm PO2 dt
where kOX(T) is given by kOX ðTÞ ¼ k0OX eðEa =Rðð1=TÞð1=T 0 ÞÞÞ , CC is the coke concentration in the catalyst (mg coke (mg catalyst)1), m is the kinetic order for coke combustion, PO2 is the oxygen partial pressure and if a temperature ramp of constant heating rate b (K min1) is used, the time–temperature relationship is T = T0 + bt, where T0 is the initial temperature. Assuming m = 1, considering 0 coke conversion as XC = (C0Ct)/C0 and k0OX ¼ PO2 k0OX ; the Table 4 Kinetic parameter values for the kinetic models chosen Parameter
Fig. 3. Correlation between experimental data (symbols) and model predictions (lines) for the temporal evolution of molar flow of propane, propene, hydrogen and cracking at the reactor exit from deactivation model D3 (mmol min1). W/ FA0 = 74 s kg mol1; ur = 1.7; propane/argon ratio = 0.5; T = 520 8C.
(12)
Value
S.D.
Propane dehydrogenation 0.5242 mmol g1 min1 bar1 k01 Ea1 34.57 kJ mol1 DH 85.817 kJ mol1 K0 3.46 k02 0.00465 mmol g1 min1 bar1 Ea2 137.31 kJ mol1 k03 0.000236 mmol g1 min1 bar1 Ea3 154.54 kJ mol1 g01 948.92 (g cat) (g coke)1 Eag1 9.61 kJ mol1 g2 399 (g cat) (g coke)1 g3 40.07
0.0.1641 9.13 22.46 0.35 0.00243 37.82 8.01E5 15.09 67.08 1.94 43.21 14.39
Coke formation k01C Ea1C k02C Ea2C Cmax
234 (mg cat) (mg coke)1 min1 38.43 kJ mol1 1.45E6 (mg coke) (mg cat)1 min1 125.51 kJ mol1 1.04E3 (mg coke) (mg cat)1
6.76 1.38 4.98E8 1.85 5.20E6
Combustion model 0 k0OX Ea
0.0357 min1 4.27 kJ mol1
0.0017 0.087
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Fig. 4. Correlation between experimental data and model prediction for the molar flow of propane at the reactor exit (mmol min1) in the range of operating conditions used.
Fig. 6. Coke conversion as a function of time in non-isothermal conditions.
The parameter values were obtained by fitting the experimental data of coke conversion. An acceptable agreement of experimental and calculated data was found (Fig. 6). Table 4 lists the parameters obtained.
kinetics data were obtained independently, as described above. In order to take into account the ascending flow of solid, a three-phase model is proposed that includes the emulsion, bubble and wake phases where the gas existing inside the bubble, including the wake, is perfectly mixed [23]. It has been assumed that the gas rises in the emulsion at the minimum fluidisation velocity and that the remaining gas rises with the bubbles. The pressure, temperature, bubble size, fraction of bed in bubbles and fraction of the bubble volume occupied by the wake were considered to be constant. Under steady-state conditions, the mass conservation equations for a given gaseous species i are the following: In the bubble and wake phase:
5. Simulation of a two-zone fluidized bed reactor
dF i;B ¼ K B;E Aða þ a f W ÞðC i;B C i;E Þ þ Ri;B rcat Aa f W ð1 emf Þ (14) dz
integrated equation has the following expression: 1 X C ¼ ð1
Z 0 X C0 Þ exp k0OX
t
exp 0
Ea 1 1 dt T 0 þ bt T r R
(13)
As stated in Section 1, one of the requirements for the kinetic model developed in this work was that it should be able to predict reasonably well the performance of a two-zone fluidized bed reactor (Fig. 7). The model used for the reactor simulation contains no adjustable parameters and is in steady-state condition. In the two-zone fluidized bed reactor propane and oxygen are fed at different levels, providing separated zones for the reaction and catalyst regeneration and, in spite of the continuous coke formation, steady state is achieved because a dynamic equilibrium between coke formation in the reducing zone and coke burning in the oxidizing zone. The
Fig. 5. Correlation between experimental data and model prediction for the molar flow of propene at the reactor exit (mmol min1) in the range of operating conditions used.
In the emulsion phase: dF i;E ¼ K E;B Að1 a a f W ÞðC i;E C i;B Þ þ Ri;E rcat Að1 a dz a f W Þ f W ð1 emf Þ
Fig. 7. Scheme of the two-zone fluidized bed reactor (TZFBR).
(15)
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The rising velocity for bubbles in a bubbling bed is calculated with the gas velocity (u0): 0:5
uB ¼ ðu0 umf Þ þ 0:711ðgdb Þ
(16)
The gas exchange coefficient between bubble and emulsion (KB,E) is calculated as [24]: ! 0:25 D0:5 umf i;gas g þ 5:85 (17) K B;E ¼ 4:5 db d1:25 b The bubble diameter is calculated as the average values given by the Mori and Wen equation [25] with the initial and final bubble diameters. The fraction of bubbles in the bed and the gas exchange coefficient between emulsion and bubble (KE,B) are calculated to assure mass balance. Regarding the flow of solid, the main variable is the coke concentration, which will vary along the bed. In this case, the mass conservation equations are as follows: In wake phase: us;W
dC C;W ¼ K W;E ðC C;W C C;E Þ þ RC;W dz
(18)
In the emulsion phase: us;E
dC C;E ¼ K E;W ðC C;W C C;E Þ þ RC;E dz
(19)
The wake velocity is assumed to be the bubble velocity, and the emulsion velocity is calculated to assure the mass balance. The exchange velocity of solid between wakes and emulsion is calculated with an exchange coefficient, using the following equation proposed by Lim et al. [26]: K W;E ¼
0:15 db
(20)
The mass balances constitute a system of ordinary differential equations at steady state. This set of equations was solved by a Runge–Kutta–Merson method, fitting the concentration of the coke in the bottom of the reactor that would close the total mass balance for all the elementary components (carbon and oxygen). A comparison between theoretical and experimental results is shown in Fig. 8. In the TZFBR, the percentage of oxygen in the feed to the reactor is one of the most important variables [5]. Fig. 8a
163
shows its effect for a set of experiments performed with a constant input of propane (306 cm3(STP) min1) and varying argon and oxygen fluxes, keeping its addition constant (306 cm3(STP) min1). It may be seen that the model predicts satisfactorily the experimental trends, although the propane conversion is slightly higher and propene selectivity slightly lower. This difference could be attributed to the assumption made in the model, where oxygen that does not react with coke is used to burn propane, and experimentally the formation of water has been observed due to the reaction of oxygen with hydrogen. Therefore, CO2 formation is overestimated and a higher propane conversion and lower selectivity to propene are predicted. This effect increases with the oxygen content. Finally, Fig. 8b shows the effect of the temperature. The simulation predicts a slight decrease of the conversion and a mild decrease of the selectivity with the temperature. 6. Conclusions A complete kinetic model of propane dehydrogenation over a Pt–Sn–K/g-Al2O3 catalyst has been developed. Extensive experimental work has been performed to obtain the kinetic parameters involved in the reaction, in the coke formation, in the activity–coke content relationship and in the catalyst regeneration. A Langmuir–Hinshelwood mechanism with strong adsorption of propene has been chosen as it provides the best fit to describe the reaction kinetics. A monolayer–multilayer coke deposition model successfully describes the coking behaviour of Pt–Sn–K/Al2O3 catalyst during the dehydrogenation of propane. The relationship between coke content in the catalyst and catalyst activity has also been studied and determined, with the activity being related with both monolayer and multilayer coke. A power-law model has been proposed to describe the catalyst regeneration. The complete model explains reasonably well the qualitative and quantitative observations presented in this work. A three-phase reactor model combined with the proposed kinetic model describes satisfactorily the conversion and propene selectivity found experimentally in the catalytic dehydrogenation of propane in a two-zone fluidized bed reactor. Moreover, it explains their variations with the operating variables. Acknowledgement The authors thank DGI (Spain) for financial support in the project CTQ2004-01721/PPQ. References
Fig. 8. (a) Influence of the % of oxygen upon propane conversion and propene selectivity in TZFBR reactor. T = 500 8C; W = 60 g; Qpropane = 306 cm3(STP) min1; Qargon+oxygen = 306 cm3(STP) min1. (b) Influence of the temperature upon propane conversion and propene selectivity. W = 35 g; Qpropane = 437.5 cm3(STP) min1; Qargon = 350 cm3(STP) min1; Qoxygen = 87.5 cm3(STP) min1.
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