Ind. Eng. Chem. Res. 1995,34, 3808-3816
3808
Conversion of Methanol to Lower Olefins. Kinetic Modeling, Reactor Simulation, and Selection A N. Renb Bos* and Peter J. J. Tromp Koninklzjke I Shell-Laboratorium, Amsterdam, P.O. Box 38000, 1030 BN Amsterdam, The Netherlands
Henk N. Akse Shell Znternationale Chemie Maatschappzj, The Hague, The Netherlands
Reactor types for commercial-scale methanol-to-olefins (MTO) processes in the ethene mode, using a small-pore molecular-sieve catalyst, have been evaluated both qualitatively and quantitatively. A kinetic model has been developed via an iterative process of model formulation, parameter estimation, and model validation. The final model consists of 12 reactions involving 6 component lumps plus coke. Important factors are the occurrence of consecutive reactions and the effect of coke on both the activity and selectivity. This kinetic model has been implemented in mathematical models of various reactors for the estimation of product selectivities and main reactor dimensions. These formed the basis for a comparison of different reactor types for a commercial-scale process. A circulating fast fluidized-bed reactor and a turbulent fluidizedbed reactor emerged as the most promising reactor systems for MTO in the ethene mode; ethene/ propene ratios of 1-1.5 can be achieved with realistic reactor dimensions. 1. Introduction The selective dehydration of methanol to olefins (MTO) is a potential route for the production of C2-C4 lower olefins. To be cost effective, the process-apart from showing a high selectivity to lower olefins-should be directed to making ethene as a main product because the price of ethene, a t present but also predicted for the future, is substantially higher than that of propene. At our laboratories, zeolite- and molecular-sieve type catalysts for MTO have been screened and optimized with respect t o selectivity, activity, and stability. The simplest and most efficient reactor configuration for testing catalysts for the MTO reaction is a fured-bed reactor. However, because the MTO reaction is highly exothermic (-50 kJ/mol) and the catalyst lifetime is limited owing to coking, in commercial practice a furedbed reactor would not be the obvious reactor of choice. The objective of the present study is to identify the best type of reactor and reactor configuration for a commercial-scale (-300 kt ethene/y) MTO process in the ethene mode. This paper describes the development of a kinetic model for a catalyst based on SAPO-34. This type of catalyst shows a high selectivity of ethene in MTO; see, for example, Inui et al. (19901, Kaiser (19851, Liang et al. (19901, and Nawaz et al. (1991). The model is based on dedicated experiments in a pulse-flow, fixed-bed reactor. The kinetic model is implemented in mathematical models of various reactors for the estimation of product selectivities and main dimensions. These form the quantitative basis for comparison of different reactor types. 2. Initial Work 2.1. Experimental Section. A catalyst screening study showed that for high ethene selectivities to be obtained, molecular sieves with a small-pore system are required as a catalyst. These scouting experiments were performed in a small fixed-bed reactor, using
* To whom correspondence should be addressed. E-mail:
[email protected].
1-500 mg catalyst (depending on the space time applied) diluted with 400 mg quartz. This bed was in contact at 1.2 bar with a continuous flow of inert gas (argon; 0.6 NVh), in which a feed pulse of methanol was injected. The feed pulse consisted of 50 vol % methanol and 50 vol % helium as an internal standard. The volume of the feed pulse corresponded, at the argon flow rate of 0.6 NVh, to a pulse length of 30 s. After contact with the diluted catalyst bed, small portions of the pulse were injected on-line into a gas chromatograph and analyzed. By use of helium as an internal standard, methanol plus dimethyl ether (DME) conversion and Cl-C6 product selectivities were calculated on a CH2 basis (CH30H CH2 HzO). In general, a mass balance of 95 wt % or better was obtained; the remaining part was attributed to the formation of coke. Independent measurements of the coke content of the spent catalyst confirmed the coke selectivities to be approximately 5%. 2.2. Initial Results. In the pulse-flow experiments described above, the product selectivity, particularly ethene selectivity, changed strongly with time-onstream; see Figure 1,which shows the result for a space velocity WHSV = 1 h-l. High ethene selectivities, up to 50 wt %, in combination with a total lower olefins yield of 90 wt % a t complete methanol conversion were obtained. For a WHSV of 1 h-l, the catalyst lifetime during these pulse-flow experiments was approximately 7 h. At higher space velocities it was shorter. One explanation for the observed increase in ethene selectivity with time-on-stream might be the enhancement of the so-called shape-selective effects by coke deposition in the cavities of the catalyst crystal structure: at increasing coke content, the free space in the cavities becomes smaller, and therefore, the formation of the larger molecules is suppressed more strongly than that of the smaller molecules (Anderson et al., 1990). At increasing time-on-stream, the coke content increases, and thus, ethene selectivity increases. However, if the coke formation is considered in more detail, there is actually a coke front that moves slowly through the catalyst bed. Direct experimental evidence for this
-
+
0888-5885f95I2634-3808~09.00f0 1995 American Chemical Society
Ind. Eng. Chem. Res., Vol. 34, No. 11, 1995 3809 Table 1. Experimental Selectivities and Conversions
Yield (%wt CH2 basis)
100
1
I
-
8 0 ~8 ~
I
Ethene
+ Propene
+ Sum butenes
Propane
++
C DME
Methanol
I ~
1
60
I
”/
1 ~
40
0
100
200
300
400
500
Time-on-stream (minutes) Figure 1. Product yields as a function of time-on-stream at a WHSV of 1 h-*.
stems from analysis of the catalyst after a limited timeon-stream: a coked upstream part and an “uncoked”-or only slightly coked-downstream part are observed. The dynamic effects associated with this moving coke front, and thus moving reaction front, severely hamper an unequivocal interpretation of these experiments, and we have concluded that these cannot form a sound basis for reactor selection and scale-up. For example, the coke distribution in a system with solids circulation differs strongly from that in a fixed-bed reactor. So, direct extrapolation of the fixed-bed reactor yield packages to industrial operation is not possible. Also, the observed increase in ethene selectivity as a function of time-onstream may be (partly) due to the decreasing contact time for all consecutive reactions as a function of timeon-stream. Furthermore, it is important to determine the main factors (time-on-stream or coke content) governing the deactivation, since these might be different for different types of reactors. The above considerations outline the limited value of the time-on-stream experiments for a proper selection and design of a commercial-scale reactor. Therefore, we have set up and performed a kinetic study of the MTO reactions under more clearly defined conditions. This kinetic study is described in the next section.
3. Kinetic Study of the MTO Reactions To reliably predict conversions and product selectivities in, for example, solids-recirculating systems (or any other type of reactor), a kinetic model is required that describes activities and product selectivities as a function of coke content and gas-phase composition (partial pressures). Since the coke content is one of the main factors governing both activity and selectivity, we decided t o perform steady state kinetic experiments in such a way that a flat coke profile existed in the bed. 3.1. Procedure for the Kinetic Experiments. Samples with varying coke contents were obtained by having a fixed-bed of fresh catalyst at 450 “C, diluted with quartz to minimize heat effects, come into contact with pulses containing 50 vol % methanol and 50 vol % helium. After a limited time-on-stream, the partially coked catalyst was removed from the reactor and separated into “coked” and “uncoked”parts. These two parts could be easily distinguished visually, the upstream coked part being much darker than the “uncoked (or rather slightly coked) downstream part. In this way, samples with approximately uniform coke contents could be obtained (3.7,8.9,and 12.3 wt %). The conversion of methanol plus dimethyl ether and product selectivities of these samples, along with those for fresh
coke space content time CH4 C2- C3-
C3
[%I
[hl
[%I
[%I [%I [%I
0 0 0 0 0 0 0 3.7 3.7 3.7 3.7 8.9 8.9 8.9 8.9 8.9 12.3 12.3 12.3 12.3 12.3
0.002 0.004 0.01 0.02 0.05 0.2
2.3 2.5 2.9 2.2 2.5 2.4 1.8 1.4 1.4 1.7 1.4 2.3 2.2 1.9 1.6 1.7 3.1 2.0 2.5 2.4 1.9
25.3 23.8 21.5 22.5 22.5 21.7 21.6 32.9 34.1 36.1 37.8 35.4 36.2 38.7 40.6 44.9 39.3 38.7 43.7 46.1 46.0
1.0
0.002 0.004 0.01 0.02 0.002 0.004 0.01
0.02 0.25 0.002 0.01 0.02 0.05 0.25
49.8 49.3 48.3 47.7 44.5 43.0 33.2 44.3 44.0 38.9 38.6 42.5 41.5 42.2 41.9 37.7 40.8 42.0 38.0 36.5 38.9
3.3 3.8 4.5 7.6 9.1 8.0
12.8 4.2 4.7
*
7.7 4.1 3.7 2.3 1.5 2.0 2.3 1.6 2.6 2.1 1.5
MeOH sum sum conversion C4[%1 C5[%1 [%I 15.0 4.3 60 15.0 5.5 76 6.7 91 15.9 15.1 4.7 99.9 16.0 5.0 100 19.9 4.7 100 22.1 7.8 100 15.2 1.9 39 14.0 1.8 61 12.2 1.4 77 11.6 2.2 97 13.9 1.9 17 13.9 2.0 23 12.2 2.7 60 11.7 2.7 67 10.4 3.2 100 12.0 6.5 13.7 2.0 33 11.9 1.3 36 10.5 1.9 63.5 9.5 2.3 100
catalyst, were subsequently measured in the fixed-bed pulse-flow reactor at 450 “C and 1.25 bar as a function of space time, covering the range of both low and high methanol conversions. The time-on-stream during these experiments was so low that no significant additional coking occurred. 3.2. Experimental Results. The results obtained with the samples containing 3.7,8.9,and 12.3 wt % coke as well as a fresh catalyst are given in Table 1. The selectivities are based on the C1-C6 products, thus excluding coke. The methane selectivities were usually approximately 2%and independent of the coke content. No significant amounts of ethane were formed. To reduce the number of components, the butanes and butenes and the C5 and C6 olefins and paraffins have been lumped into a C4 fraction and a C5 fraction, respectively. For the fresh sample, the olefinicity of the C4 and C5 fractions is about 80%; this increased with increasing coke content. The results clearly show that the activity, i.e., overall methanol conversion rate, decreased whereas the ethene selectivity increased substantially with increasing coke content of the catalyst. The selectivities to the C3-C5 fractions decreased with increasing coke content. Propene selectivity also decreased significantly a t increasing space time. For a fresh sample, propene selectivity decreased at increasing space time even a t 100% conversion. With the exception of the fresh sample where ethene selectivity is about constant, ethene selectivity increased with space time. 3.3. Model Formulation. We strived for a kinetic model that was as simple as possible; it was not our goal to unravel all details of the MTO reactions but merely to describe the main effects and trends with sufficient accuracy t o enable a comparison between alternative reactors and reactor configurations. Our first choice concerned the number of components and component lumps, i.e., the components methane (Cl), ethene (C2=), propene (C3=), and propane (C3), the lumps of butenes and butanes (C4), the lump of all heavier components lumped as C5 (pentenes being the main fraction of the latter), and coke. Our initial, simplest, reaction scheme consisted of seven parallel reactions of methanol to the six component lumps plus coke and one consecutive reaction of propene to ethene. At this stage, it is important t o distinguish between the reaction scheme and the mechanism of the reactions occurring on the catalyst surface. The latter most likely
3810 Ind. Eng. Chem. Res., Vol. 34, No. 11, 1995 consists of a complex set of consecutive reactions of the surface species (Dah1 and Kolboe, 1993), for example, adsorbed methanol gives adsorbed ethene which reacts further to adsorbed propene, etc. However, with respect to the gas-phase species this may well correspond to primarily parallel reactions. Since our conversion vs space time experiments indicated overall first-order behavior in methanol, the seven primary reactions were assumed t o be first-order in methanol Ti
= k$M,oHP
i = 1-7
(1)
=k 9 c 3 2
(2)
All rate constants are assumed to be dependent on the coke content of the catalyst. We have tried different, empirical correlations:
c being the weight percentage of coke on the catalyst. Note that the effect of coke on the selectivities, i.e., the ratio of the rates of reactions, can now be modeled by taking different values for the empirical constants ai. The larger the value of ai the more strongly the reaction rate ri decreases with increasing coke content. Mechanistically, on the basis of the shape-selectivity effects, it is expected that ai increases with increasing characteristic molecular dimensions of the corresponding components (-carbon number). For the evaluation of the kinetic experiments the pulse-flow reactor was modeled as follows. First we define weight fractions, yields, and selectivities on “a dry basis”, i.e., on the basis of CH2 equivalents. For a feed with no hydrocarbons other than methanol, it is easy t o derive W$eOH
=
WMeOH
MeOH
kc =
MeOH
On introduction of the variable space time defined as
--
1
Pcatr
[MeOHl3600 32
10-3 .
.
n
-
Pcat
3.6 x 32 When the rate constant k is expressed per bar (rather than per Pa) this yields for T = 450 “C and Peat = 800 kg/m3
&k = 0.42
and the consecutive reaction first-order in propene r8
conventional units using
[Pdbarl or kc = 0.42k
(12)
3.4. Evaluation and Discrimination of Alternative Kinetic Models. The kinetic and reactor model was implemented in SimuSolv code. SimuSolv is a commercial program both for simulation and for parameter estimation, using a sophisticated maximum likelihood approach (Dow Chemical Company, 1993). The four data sets comprising conversion and selectivity as functions of the space time with coke contents of 0,3.7,8.9, and 12.3 wt %, respectively, have first been fitted separately. This yielded the k-values of each model for each coke content. To discriminate between the alternative empirical correlations of eqs 3-6, we have plotted these fitted values as a function of the coke content. This showed eq 6 (exponential dependency)to be the best representation for the effect of coke on the rate of reactions. This procedure also yielded first estimates for the parameter values ai. Subsequently, all four data sets were fitted simultaneously, yielding values for all rate constants Kp and all ai’s. The values found for the coke parameters ai were in line with the expectation on the basis of the shape-selectivity effects: a increases with increasing carbon number. It soon appeared that the simple model of seven parallel reactions and one consecutive reaction is not able to describe the main effects. The reaction of propene to ethene does take place but only to a minor extent. This was verified by a separate experiment using propene as feed gas; some ethene but mainly butene was formed. With ethene as a feed no reaction was observed. On the basis of the observed dependencies of the selectivities on space time, we added the formation of propane from propene to the reaction scheme (reaction 91, to describe the increase in propane with increasing space time. In addition, we added the formation of butenes from propene (reaction 10) and coke from propene (reaction 11). All these consecutive reactions were assumed to be first-order in propene:
and assuming plug-flow behavior, the mass balance for ethene for our initial simple model can be written as
where (10) The equations for the other components are similar. All first-order rate constants k may be translated into
This model could predict most of the trends, but generally the predicted dependencies of the selectivities on the space time were much less pronounced than the observed ones. A major improvement was obtained by assuming that reaction 8, the formation of ethene from propene, is not only first-order in the hydrocarbon concentration but also first-order in methanol:
Ind. Eng. Chem. Res., Vol. 34, No. 11, 1995 3811 MEOH
= kE%!3~xMeOHp
(17)
= kgc3-P
(18)
'10
=K1dcC32
(19)
rll
= K11xC39
(20)
methane r9
sum C4
J
sum C5 coke
Figure 2. Final reaction scheme from model discrimination.
This implies that for large space times, i.e., complete methanol conversion, this reaction ceases to proceed, contrary t o the other consecutive reactions. Direct support for this was given by the experiments using the fresh catalyst (0% coke), where ethene selectivity remains virtually constant a t increasing space time from 0.02 to 1 h, whereas clearly propane and butenes are formed from propene. Note that the methanol conversion at a space time of 0.02 h is already 99.9%,so at space times higher than 0.02 h, virtually no methanol is present. A mechanistic justification for assuming the propene-to-ethene reaction to be first-order in both propene and methanol might be given on the basis of a chain-growing mechanism (Dahl and Kolboe, 1993) on the catalyst surface; the CH2 group of methanol is inserted into an adsorbed propene species, which subsequently cracks into ethene, so both methanol and propene are required for this consecutive reaction. In case coke is deposited on the catalyst, the dependencies on the space time are different, e.g., ethene selectivity increases significantly at increasing space time, also at virtually complete methanol conversion. To account for this, we added reactions 12, the formation of ethene from butene, and 13, the formation of coke from ethene, to our scheme. Note that by means of the coke parameters a, we can account for a difference in significance of these reactions for a fresh (low coke content) catalyst and for coked catalysts; by taking a relatively high value for the aL,the corresponding reaction decreases more strongly with increasing coke content as compared to reactions with a lower value of a. The occurrence of reaction 12, the formation of ethene from butene, was verified by a separate experiment using butene as feed gas, which yielded ethene and also propene (over a fresh catalyst). Second-order behavior proved to be more appropriate than first-order. For this model, the optimization procedure of SimuSolv converged to two very different solutions, with parameter values corresponding to either relatively low or high rates of reaction for propene t o butene (ll),butene to ethene (121, and ethene t o coke (13). In the latter case, the value of a for reaction 13 was also very high, corresponding t o a sharp decrease in its rate with increasing coke content. However, the coke selectivity, as predicted in that case, would be of the order of 30% instead of the experimentally obtained 4 4 % . The results of the experiment using butene as feed gas did not agree with the model predictions; there was no extreme coke formation and the rate of reaction was much lower than predicted by this model. Therefore, reaction 13 was deleted from the model, i.e., k 1 3 was set t o zero, t o force the program to converge t o the first solution, with relatively low rates of reactions 11 and 12. This finally leads to the model of Figure 2 and eqs 1, 6, and 17-21:
r12
= kl!$C4pxMeOHp
(21)
The parameter values and standard deviations as determined using SimuSolv are given in Table 2. The comparison between the experimental results and our final model predictions is given in Figure 3. Although there certainly remain differences between the observed and predicted selectivities, the main effects are described with sufficient accuracy to allow a comparison between alternative reactor types. The reaction rates, and thus the rate constants, are based on pure catalyst, expressed per unit mass of catalyst. For industrial application, a binder material will be added to the zeolite crystals, the catalyst typically being only 25 wt % of the total. For the evaluation in the next section, we have assumed a dilution of catalyst with binder by a factor of 4. This can be easily accounted for; the rates and rate constants expressed per unit mass of solids are smaller by a factor of 4. Note, however, that this simple approach will fail for too large particles. In that case mass transfer resistances in the meso- and macropores will not remain negligible. A rough estimation indicated these resistances to become significant for particle sizes of approximately 0.2-0.4 mm. 4. Evaluation of Reactors for MTO (Ethene Mode)
4.1. Selection Criteria. On the basis of the kinetic study, the main aspects governing the selectivity and activity of the MTO process may be summarized as follows. Coke content of the catalyst: At increasing coke content the selectivity to ethene increases. The coke content is the main factor determining the selectivity. The effect of coke is most pronounced at low coke contents. Occurrence of consecutive reactions: Consecutive reactions cannot be neglected. The net effect of these is an increase of the ethene selectivity with increasing space time (except for a fresh catalyst, where virtually no effect is found). This implies that deviations from plug-flow of the gas-phase increase the ethene selectivity at the cost of activity. Operation at too low space velocities, i.e., oversizing of the reactor, has a positive effect on the ethene yield, except for a fresh catalyst for which a net effect is virtually absent. Rate of MTO reactions and coke formation (deactivation): The MTO reaction on a fresh catalyst is very fast, with an overall first-order rate constant of roughly 250 11s. The activity decreases exponentially as a function of the coke content (and not of time); at 10% coke, the activity roughly decreases by a factor of 7. The rate of coke formation is about 4% of the rate of the MTO reaction, independent of the coke content. Other aspects that need t o be considered are as follows. Large heat effect of the MTO reaction: The adiabatic temperature rise for a feed of pure methanol is of the order of 800 K, in a solids-circulating system a majorpart of the reaction enthalpy may be taken up by the
3812 Ind. Eng. Chem. Res., Vol. 34, No. 11,1995 selectivitv (%wt. CH2 basis)
selectivitv (%wt. CH2 basis)
10 7
0
3
0,005 0,Ol 0,OZ 0,05 0,l space time (kg.h/kg)
0,2
ethene
............
................. ....
.............
Y
n
U
-
$001 0,002
. .-'-
1
40
sum C4
propene
0,5
1
0' 0,001 0,002 ethene
propane
A
0
0,005
............
-
.................................... e
Q -
I
0,Ol 0,02 0,05 space time (kg.h/kg) s u m C4
propene
A
-- - -
0,l
0,2
0,5
propane
0
-
selectivity (%wt, CH2 basis)
401
........................
c
10
1
l
I
i
1
e
30 20
.
i
30
i
I
7
20
A C
t
P
$06
~
I
1
0,002 0,003
ethene
............
propene
e
0,005 0,Ol space time (kg.h/kg) sum C4 A
0,02 propane
-
0,03
0,05
.
ethene
-
-9-
4
0 0,001 0,002
0,005 propene
0,Ol 0,02 0,05 space time (kg.h/kg)
_____
s u m C4 A
0,l
_ _ _ _ _ propane-
0,2
03
-
Figure 3. Comparison between experiments and medictions of final model: (A) 0 wt % coke; (B) 3.7 wt % coke; (C) 8.9 wt % coke; (D) 12.3 wt % coke.
Table 2. Fitted Parameter Values for the Kinetic Model of Figure 2 and Eqs 1,6,and 17-21 reaction 1 2 3 4 5 6 7 8 9 10 11 12
rate constant standard ko [ l h barla,b deviation 12.4 149.5 297.2 33.3 103.8 34.2 30 31.8 0.26 0.04 0.31 260
1.75 17.6 37.9 4.8 16.2 4.32 42 31 0.11 0.3 0.15 154
parameter
a [l/wt %Ib 0.21 0.17 0.227 0.292 0.23 0.35 0.3 0.12 0.3 0.3 0.35 0.12
standard deviation 0.012 0.012 0.011 0.014 0.012 0.5 C
0.05 0.5 C C
0.05
See eqs 9-12; for second-order reactions l/(h bar2). See eq 6. Not available.
solids, reducing the adiabatic temperature rise within the reactor. An overall heat balance, however, excluding the regenerator, shows that the heat of reaction (49 kJ/mol) can approximately be taken up by the heat of vaporization of methanol (35 kJ/mol) and the heat capacity of the feed stream (starting with liquid methanol at ambient conditions). Detrimental effect of increased particle size on the activity: The high rate of reaction indicates that particles of, say, 1mm are severely influenced by diffision resistances in the macropores; so for fixed- or movingbed type reactors, a compromise must be found between activity and pressure drop. Considerations as "well known and proven technology" have also been taken into account. We conclude that in order to achieve high ethene selectivities, the reactor system should ideally be such that the main fraction of the methanol is converted on a coked catalyst and that contact between methanol feed
and fresh catalyst (high activity, low selectivity to ethene) is avoided. This is best achieved in a system where gas and solids are contacted countercurrently. We have selected a fixed-bed reactor (swing operation), a riser reactor, a countercurrent moving-bed reactor, gas-solid trickle-flow reactor, a circulating fast fluidized-bed reactor, and a turbulent fluidized-bed (with and without staging of the solids). In these reactors our main objectives might in principle be achieved. The purpose of the modeling presented below is to obtain a quantitative comparison between the various selected reactor types and to evaluate the influence of the process parameters on the selectivities and main dimensionsof these reactors. The regenerator for decoking of the catalyst from the reactor will not be considered in this study. All calculations were made for a pressure of 1.7 bar, a temperature of 450 "C, and an ethene production of about 300 kt/y, with an overall methanol conversion of 99%. 4.2. Dynamic Fixed-Bed Reactor. For a fixed-bed operation-and for the description of the time-on-stream experiments in the pulse-flow reactor-we need t o account for the formation of coke as a function of both time and space. Realizing that the time constant for the buildup of coke on the catalyst is considerably lower than the time constant for conversion of the gaseous components, we may apply the pseudo-steady state approximation for the gas-phase mass balances, i.e., the same balances as described before can be used, albeit that the coke concentration on the catalyst is not constant but varies both in time and in space. This latter is described by the following differential equation:
(22) Note that we now have a set of partial differential
Ind. Eng. Chem. Res., Vol. 34,No. 11, 1995 3813 equations. The detailed description of this model and its application t o evaluate fmed-bed reactors in swing operation are not included in this paper since this reactor type does not appear to be a realistic option. We have rejected the use of fxed-bed reactors in swing operation as a feasible option for a commercialscale MTO process because of the following. (1) Due to the high exothermicity, an expensive cooled multitubular reactor would be required. (2) Several such reactors would be needed because of the rapid deactivation rate. Time-on-stream periods of more than 1 h are only feasible by strongly oversizing, i.e., using space velocities of 5 h-l or lower, whereas, for a fresh catalyst, a space velocity of 50 h-' may be applied to achieve 99% conversion. (3) Regeneration is more complex compared to a solids circulation system. (4) The pressure drop of such reactors, assuming particle sizes of 0.4mm, would be excessively high. For larger particles, with a significantly lower pressure drop, the apparent activity would be much lower. We show in parts A-D of Figure 4 the results of a simulation of the first 2 h of the time-on-stream experiments, which clearly illustrate the slow movement of the coke front through the bed (Figure 4A). Figure 4B shows the (initially fast) increase of ethene selectivity, whereas the propene selectivity (Figure 4C)a t the outlet remains roughly constant as function of time-on-stream, as was also observed experimentally. This is due to, on the one hand, an increase of the coke content in the main reaction zone and, on the other hand, the decreasing length of the catalyst bed beyond this main zone. 4.3. Riser Reactor. The riser in a conventional riser-regenerator configuration is modeled using the following assumptions: gas is in plug-flow; solids are in plug-flow; gas and solid velocities are equal (no-slip). On the basis of these assumptions, we can describe the axial coke profile by means of de- - 100-l4
dz*
(k,w$eoHP* 32 CATOIL
+ ko,w&+€*')
(23)
It should be noted that it is questionable whether or not incomplete regeneration, required to obtain a nonzero coke content at the riser inlet, can indeed be achieved. Furthermore, it is unclear whether a partially coked catalyst from a reactor is comparable to a partially decoked catalyst from a regenerator. The potential impact of this latter point should not be underestimated, and further research is required if this option is to be pursued. 4.4. Moving-BedReactor and Gas-Solid TrickleFlow Reactor. We have also developed a model for a moving-bed type reactor and a gas-solid trickle-flow reactor. Note that in both types countercurrent gassolids contact can be achieved, the main differences between these two being the catalyst holdup and, consequently, the pressure drop. The same model can be used for both reactors. We have discarded the moving-bed reactor as a feasible option basically for the same reasons as those described for fixed-bed reactors: the large exothermicity (in a moving-bed the solids take up only a minor fraction of the reaction enthalpy, except at unrealistically high solids flow rates) and also the pressure drop would be excessive. In a gas-solid trickle-flow reactor these two problems may be alleviated to some extent. However, this type of reactor may not be regarded as proven technology yet. Therefore, we did not evaluate these two reactor types in more detail. 4.5. Fluidized-BedReactors. 4.5.1. Circulating Fast Fluidized-BedReactor. A circulating fast fluidized-bed reactor consists of a (dense phase) riser reaction section and a recycle section in which heat exchange takes place. A high recirculation ratio reduces the temperature rise in the riser section. Note that this does not refer to the circulation between reactor and regenerator; as mentioned before, the regenerator is not taken into account in this study. If we define zcat as the total residence time in the reactor part (riser) of the system, we may write for the residence time per pass
which must be solved simultaneously with all mass balances for the gaseous components:
(26)
(24)
R being the recirculation ratio. Using an assumption
The "catoil" ratio is defined as kilogram of catalyst per kilogram of methanol. By use of the no-slip assumption, it is easy to derive
about the solids-gas slip velocity in the riser part, e.g., the solids velocity is 0.7 times the gas velocity (Kunii and Levenspiel, 19911, we can relate reatto the recirculation ratio
P W~,~,CATOIG €cat
-
Pcat
+
1 wM~OHCATOILPP
(27)
(25)
R
Pcat
A solids holdup of around 2%is a realistic value for FCC type riser operation. Therefore, the corresponding value of 30 for the catoil ratio was used for evaluation of the effect of the other parameters. A gas velocity of 10 d s was chosen as a base case. Figure 5 shows the effect of inlet coke content on the product selectivity and riser length for the specified base case conditions. This shows the compromise between the reactor length and the ethene/propene ratio. The calculations clearly show that ethene-to-propene ratios of around 1 cannot be achieved with a completely regenerated catalyst, whereas for the partially regenerated catalyst, with acceptable ethene selectivity, the riser lengths become quite high.
+ 1= 0
€cat Pcat 1 . 7 ~ 1 - ccat pg CATOIL
(28)
in which CATOIL =
€cat
Vi-Pcat
(29)
zcat$MeOH
The adiabatic temperature rise in the riser section, follows from
3814 Ind. Eng. Chem. Res., Vol. 34, No. 11, 1995
A
0
0.0
0.2
0.6
0.9
1.0
0.8
0.0
0.2
0.4
SPACE TIME (hl
0.6
0.8
1.0
SPACE TIME [h]
s 0
E
t t
[ 1 g
0
0.2
0.0
0. ri
0.6
0
1.0
0.8
0.0
SPACE TIME [h]
0.2
0.u
0.6
0.8
1.0
SPACE TIME Ihl
Figure 4. Simulation of the first 2 h of the time-on-stream experiment of Figure 1 (note that space time corresponds to the length coordinate): 0, t = h; 0, t = 0.4 h A t = 0.8 h; t = 1.2h; x , t = 1.6 h; 0 , t = 2 h; A, coke profile; B, ethene selectivity; C, propene selectivity; D, C4 selectivity.
+,
’ eth_ene
propene s
e reactoLlength 1 ;E
-1 50
I100
25
50
0‘ 0
the coke content to the residence time. In our kinetic model, the coke-formation rate depends on the methanol partial pressure-and to a minor extent also on the propene partial pressure-which is not constant throughout the reactor. Therefore, we assume an average methanol partial pressure equal to the logarithmic mean of the partial pressures at the inlet and outlet. In that case, neglecting the minor effect of coking from propene (reaction 101,
1
1
2
3 4 5 inlet coke content (%wt)
6
7
with
Figure 5. Effect of inlet coke content on product selectivity and riser reactor length.
We assume the gas phase to be in plug-flow and the solids to be ideally mixed. The main variable governing the selectivity and activity is the coke content of the catalyst. Therefore, our main concern is a proper description of the coke distribution and its effect on the overall selectivities and activity. In the case of ideal mixing of the solids we have
E(t)= Le-t/%,t
yielding
The mean coke content can then be evaluated from
(31)
rc*t
To calculate the coke distribution corresponding to this solids-residence-timedistribution, we need to relate
However, it would not be correct to use this average coke content to determine “average”rate constants. This
selectivity (w%, CH2-basis) 100
bed volume (m**3)
eth_ene
p E e
s
. bed.v$!ume 800
80
600
60
60
400
40 20
20
200 a...'a
0
0
5
10
15
20 25 30 35 40 solids res. time (min )
45
50
55
Figure 6. Effect of solids residence time on product selectivity and bed volume for a one-stage fast fluidized-bed reactor.
would only be allowed if a, were the same for all reactions. Since it is the difference in these values that largely reflects the effect of coke on the selectivity, this should be accounted for to properly describe the effect of a coke distribution. A better approximation is obtained by means of "average" rate constants defined by
SL= h%,(c(t))E(t)dt
oi
60
5
10
15
20 25 30 35 40 solids res. time (min)
6
50
55
0 6'0
Figure 7. Effect of solids residence time on product selectivity and bed volume for a turbulent fluidized-bed reactor. ethene selectivity (%wt, CH2 basis)
(36)
Substitution of eqs 31 and 34 yields 3"
(37)
0
~
2
4
6 8 10 12 mean coke content (%wt)
14
16
18
Figure 8. Effect of staging in circulating fluidized-bed on ethene
Equation 37 was solved both analytically and numerically. The latter method proved to be more convenient to implement within the SimuSolv program. To suppress contact between fresh catalyst and methanol feed, staging of the solids is preferred. To this end, two circulating fast fluidized-bed reactors can be placed in series (which may actually be two circulating sections within one vessel). This will be referred to as a two-stage circulating fast fluidized-bed reactor. Using the ideal mixers in series approach for a two-stage fluidized-bed, we have calculated the rate constants with
ko xi = -Jp + a$, t> M
rcat
-ai/a7t
e-t/rCat dt
(38)
rcat
z,,t being the solids residence time per stage. As a base case, we assumed a solids holdup of lo%, which is a typical value for fast fluidization (Kunii and Levenspiel, 1991). At lower holdups, the bed length increases (inversely proportional). The base case gas velocity was 3 d s . Figure 6 shows the effect of solids residence time on the selectivities and bed volume. The residence time refers t o the total solids residence time within the reaction zone of the circulating bed, thus the residence time per pass multiplied by R + 1. It can be seen that for a solids residence time of roughly 7 min, the coke distribution is such that a n ethene/propene ratio of 1 is achieved, the average coke content being 8.2 wt %. At increasing residence time, ethene selectivity increased a t the cost of a larger reactor. 4.5.2. Turbulent Fluidized-BedReactor. Figure 7 shows the results of calculations for a turbulent fluidized-bed reactor, which from a modeling point of view is similar to the circulating fluidized-bed reactor. As compared to a fast fluidized-bed reactor, in a turbulent fluidized-bed the gas velocities are lower and
selectivity.
a higher solids holdup is achieved. We have assumed a solids holdup of 25%. Obviously, the only effect is a decrease in reactor volume for a given residence time. As a base case, the gas velocity assumed was 1 d s , which leads to a diameter of 8.5 m for the required throughput. The simulation showed that ethene/propene weight ratios larger than 1 can be obtained with realistic values of the reactor dimensions. The positive effect of staging on the selectivity can be seen by comparison of the ethene/propene ratios at the same total solids residence time or, more conveniently, a t the same average coke content (see Figure 8). For two stages, the selectivity to ethene is considerably higher, which is of course due to the fact that for two stages the coke distribution is narrower. The fraction of catalyst with a low coke content (and thus a relatively active catalyst and a low ethene selectivity) is much lower in the case of two stages as compared to that in one ideally mixed stage. At the same average coke content, the two-stage reactor yields a higher selectivity to ethene at a somewhat larger bed volume. 4.6. Final Selection. In the previous paragraphs, we discussed simulations of several reactor types and configurations. As indicated before, a compromise must be sought between reactor size and ethene yield. This trade-off between additional margin and additional capital investment can only be made on the basis of an economic objective function. The results of this economic optimization fall beyond the scope of this publication. 6. Conclusions
The kinetic study shows that the MTO reactions on a fresh catalyst are very fast, with an overall first-order
3816 Ind. Eng. Chem. Res., Vol. 34, No. 11, 1995
rate constant of roughly 250 mlasm&-'. The coke content of the catalyst is the main factor governing the selectivity and activity of the catalyst. In order to achieve an ethene-to-propene ratio of 1 or higher, a t least 7-8 wt % of coke must be present on the catalyst. Consecutive reactions cannot be neglected. The net effects of these are an increase of ethene and propane selectivity and a decrease of mainly propene selectivity. By use of the SimuSolv modeling and parameter estimation program, several kinetic models have been evaluated. An initially proposed simple scheme of seven parallel reactions and one consecutive reaction was not able to describe the main effects. The final kinetic network of 10 first-order and 2 second-order reactions describes the experimental results satisfactory. The effect of coke on the selectivity and activity could best be described by an exponential dependency of all reaction rate constants on the wt % of coke on the catalyst. The different values found for the empirical decay constants for each reaction properly reflect the shape-selective effects induced by coke deposition. From the comparison of different reactor types selected for further evaluation, viz., fixed-bed reactor in swing operation, moving-bed reactor, conventional riser, turbulent fluidized-bed reactor, and circulating fast fluidized-bed reactor, with and without staging of the solids, the circulating fast fluidized-bed reactor and a turbulent fluidized-bed reactor emerge as the most promising reactor systems for MTO in the ethene mode. Ethene/propene weight ratios larger than 1 can be obtained with realistic values of the reactor dimensions for a turbulent fluidized-bed reactor. A compromise must be found between reactor size and ethene yield, which is only possible via an economic optimization study; see Akse et al. (1994).
List of Symbols c = weight percentage of coke on the catalyst, kg/(lOOkgcat)CATOIL = ratio of catalyst mass flow rate to methanol mass flow rate, kgca&gMeOH C, = heat capacity, J/(kg K) E ( t ) = residence time distribution function, l/h AHr = reaction enthalpy, J/mol k o = rate constant for zero coke concentration, l/(h bar") k, = rate constant of reaction i, l/(h bar") k , = conventional first-order rate constant, mla$(mza, s) k y = modified rate constant defined by eq 32, l/h [MeOH] = methanol concentration, mol/m3 M = molar mass, kg/mol n = reaction order P = pressure, bar P* = methanol inlet partial pressure, bar r, = rate of reaction i, kgreactant/(kgcath) r:= rate of reaction i, based on CH2, kgcaz/(kgcath) R = universal gas constant, J/(mol K) R = recirculation ratio t = time, h T = temperature, K ATad = adiabatic temperature rise, K u = superficial velocity, m / s V = volume, m3 w, = weight fraction of component j , kg,/kg wJ" = weight fraction of component j on a CH2 basis,
x, = molar fraction of component j , mol/mol a, = empirical deactivation constant defined by eq 6 ,
1OOkgcaJkg ccat = solids holdup, rnzat/m3 P b = catalyst bulk density, kgcaJrn3 peat =
catalyst particle density, kgcat/m;?art,cle
pg = gas density, kg/m3 #,, = mass flow rate, kg/h t* = space time, (kgcath)/kgM,O, teat = catalyst residence time, h tpass = residence time per pass, h ( = relative degree of methanol conversion
Sub- and Superscripts -
= mean value cat = catalyst g = gas phase in1 = inlet conditions r = reactor
List of Abbreviations C1 = methane C2= = ethene C3= = propene C3 = propane C4 = lump of butenes and butanes C5 = lump of components with more than four carbon atoms CH2, = CH2 group of component i DME = dimethyl ether MeOH = methanol
Literature Cited Akse, H. N.; Bos, A. N. R.; Tromp, P. J. J. Etheen uit Methanol? NPT Procestechnol. 1994, 11, 25-29 (in Dutch). Anderson, M. W.; Sulikowski, B.; Barrie, P. J.; Klinowski, J. In Situ Solid-state NMR Studies of the Catalytic Conversion of Methanol on the Molecular Sieve SAPO-34. J . Phys. Chem. 1990,94,2730-2734. Dahl, I. M.; Kolboe, S.On the Reaction Mechanism for Propene Formation in the MTO Reaction over SAPO-34. Catal. Lett. 1993,20, 329-336. Dow Chemical Company. SimuSolv Modeling and Simulation Software, version 3.0; Dow Chemical Company: Midland, MI, 1993. Inui, T.; Matsuda, H.; Okaniwa, H.; Miyamoto, A. Preparation of Silico-Alumino-Phosphates by the Rapid Crystallization Method and their Catalytic Performance in the Conversion of Methanol to Light Olefins. Appl. Catal. 1990, 58, 155-163. Kaiser, S.W. Methanol Conversion to Light Olefins over Silicoaluminophosphate Molecular Sieves. Arabian J . Sei. Eng. 1985, 10 (41, 361-366. Kunii, D.; Levenspiel, 0. High Velocity Fluidization. In Fluidization Engineering, 2nd ed.; Butterworth-Heinemann: Woburn, MA, 1991; pp 193-210. Liang, J.; Li, H.; Zhao, S.; Guo, W.; Wang, R.; Ying, M. Characteristics and Performance of SAPO-34 Catalyst for Methanolto-Olefin Conversion. Appl. Catal. 1990, 64, 31-40. Nawaz, S.; Kolboe, S.; Kvisle, S.; Lillerud, K.-P.; Stocker, M.; Oren, H. M. Selectivity and Deactivation Profiles of Zeolite Type Materials in the MTO Process. In Natural Gas Conversion; Holmen, A,, Ed.; Elsevier Science Publishers B.V.: Amsterdam, The Netherlands, 1991; pp 421-427.
Received for review September 7, 1994 Accepted April 5, 1995@
IE940528A
~~CHZ&~CHP
WHSV = weight hourly space velocity (on a methanol basis), kgMeod(kgcat h)
@
Abstract published i n Advance A C S Abstracts, October 15,
1995.
