MTG Process in a Fixed-Bed Reactor. Operation and Simulation of a

Ind. Eng. Chem. Res. 2001, 40, 6087-6098

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MTG Process in a Fixed-Bed Reactor. Operation and Simulation of a Pseudoadiabatic Experimental Unit Andre´ s T. Aguayo,* Ana G. Gayubo, Marta Castilla, Jose´ M. Arandes, Martin Olazar, and Javier Bilbao Departamento de Ingenierıá Quı´mica, Universidad del Paı´s Vasco, Apartado 644, 48080 Bilbao, Spain

A simulation model has been proposed and used to predict the behavior of a laboratory pseudoadiabatic reactor for the transformation of methanol into hydrocarbons. The model incorporates recent advances in the kinetic modeling of the main reaction steps and of the deactivation by coke of the catalyst in the 350-450 °C range and is efficient for analyzing the effect of process conditions (space time, water content and inert gas in the feed, and inlet temperature) on the evolution with time on stream of the temperature profile along the reactor and on the yields of the product lumps. Introduction The design of a reactor to be used for the MTG (methanol-to-gasoline) process is limited by the high exothermicity of the reaction and by the rapid deactivation by coke of the catalyst (prepared based on HZSM-5 zeolite). The aim of this paper is to update the design of a fixed-bed reactor for this process by incorporating recent advances in the kinetic modeling of the main reaction1,2 and of the deactivation.3-5 The simulation results have been verified by operation of the reactor in a laboratory unit. The information obtained is complementary to that in the literature on an industrial operation in New Zealand.6-9 A noteworthy precedent is the design study of an original fixed-bed reactor with external cooling,10-13 although in that case, catalyst deactivation (the main problem in the MTG process) was not taken into account. A considerable number of design studies for reactors subjected to deactivation and operation in the adiabatic regime correspond to metallic catalysts. In these studies, when the cause of deactivation is the adsorption of poisons, relatively simple kinetic models can be used because deactivation is independent of the concentration of the main reaction components.14-18 In reactions in which the catalyst deactivates by coke deposition, certain authors have studied the design of reactors by using potential kinetic models.19,20 Other authors have usedLangmuir-Hinshelwood-Hougen-Watson(LHHW) kinetic models for the main reaction and for the deactivation.21-23 At the same time, experimental techniques have been described in the literature for verification of the profiles calculated for the temperature and for coke deposition along the reactor.24-27 In processes on acid catalysts, such as the MTG process, knowledge of the evolution with the time on stream of the temperature profile along the reactor is very important because the acid sites are very sensitive to temperature. In addition to better knowledge of the acid sites of HZSM-5 zeolite and of their role in the reaction mechanism, important advances have also been made in the understanding of the limitations of the hydrothermal stability of this catalyst and of the problems involving irreversible deactivation by dealumination.28,29 This effect is very important because it * Corresponding author. Tel.: 34-94-6012580. Fax: 34-944648500. E-mail: [email protected].

restricts the operation range (temperature and water content) of the reaction. The design of a reactor for the MTG process is limited by the complexity of the reaction scheme. After the industrial introduction of the MTG process, significant advances have been reported in the literature in topics of great importance in the design of reactors for this process. In addition to advances in the understanding of the reaction mechanism,9,30-32 kinetic schemes with lumps and kinetic equations for the reaction at zero time on stream (fresh catalyst)33-36 and for the deactivation by coke have been proposed.3,4 Recently, the effect of the water content in the reaction medium has been quantified in the kinetic modeling of the main reaction and of the deactivation.2,5 Water simultaneously attenuates the individual steps of the kinetic scheme and the deactivation, because it competes in the adsorption on the active sites, on one hand, with oxonium ions and other intermediate components of the main reaction and, on the other hand, with coke precursor components. Experimental Section Equipment. The reaction equipment used (Figure 1) has already been described.37 The reactor (Figure 2) is made of 316 stainless steel and is 0.028 m in internal diameter, 0.030 m in external diameter, and 1 m in length. The catalyst bed is located on a distributor plate with orifices of 1.5-µm diameter. The zone for preheating the feed is 0.032 m in internal diameter and 0.235 m in length. The reactor is surrounded by a stainless steel encasing containing six electric resistances placed in succession. The first one is of 1000 W and is located in the preheating zone. The other five are of 750 W each. The role of these resistances, which are individually controlled, is to maintain a pseudoadiabatic regime by compensating for the heat loss across the external wall of the encasing. Before the reaction, using a stream of inert gas, the power required to compensate for the heat losses in each one of the resistances is established, so that a uniform temperature is reached. This power is maintained throughout the runs. In the feed preheating zone is a device for rapidly correcting the reactor inlet temperature. It consists of a tube with pressurized air as the cooling fluid that is activated when the temperature is 2 K above the set point. The reactor is provided with type-K thermo-

10.1021/ie0101893 CCC: $20.00 © 2001 American Chemical Society Published on Web 12/04/2001

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Ind. Eng. Chem. Res., Vol. 40, No. 26, 2001

Figure 1. Reaction equipment.

Figure 2. Reactor scheme.

couples that are 0.05 m apart from one another. At each longitudinal position, there are three thermocouples for measuring the temperature at three radial positions in the bed (axis, intermediate zone, and wall). The products

pass through a 10-port valve that allows for a sample to be sent to a Hewlett-Packard 5890 Series II chromatograph. The feed-reaction-analysis system is controlled by a computer operating with a program in Fortran. Product separation is carried out using an arrangement of three columns: (1) an HP-1 semicapillary of 0.53-mm diameter and 5.0-m length, which splits the sample into two fractions, namely, (a) volatile (C1-C4) and polar (methanol, water, and DME) components and (b) the remaining products; (2) a SUPEL-Q Plot semicapillary of 0.53-mm diameter and 30.0-m length for complete separation of volatile and polar components, which will be analyzed using TCD (thermal conductivity detection) and FID (flame ionization detection); and (3) a PONA capillary of 0.20-mm diameter and 50.0-m length for separation of the remaining products, which will be analyzed by FID. The detailed identification of the components analyzed by chromatography was performed by means of GC-FTIR spectroscopy (Hewlett-Packard 5890 IINicolet 740) and of GC-MS (Hewlett-Packard 5890 IIEngine). Identification of the components (isoparaffins, olefins, aromatics, and naphtenes) was done using PIONA standard samples for chromatography from Air Liquide. Calibration of the areas corresponding to the chromatographic peaks was done using specific factors for each component. These factors were compared using commercial mixtures from Air Liquide and methanolwater mixtures at given proportions. Catalyst. The ZSM-5 zeolite was synthesized with a Si/Al ratio of 24 from sodium silicate, aluminum sulfate, and tetra-n-propylammonium bromide and was cationexchanged to HZSM-5 with ammonium nitrate. The method used has been detailed in a previous paper38

Ind. Eng. Chem. Res., Vol. 40, No. 26, 2001 6089 Table 1. Properties of the HZSM-5 Zeolite and of the Catalyst HZSM-5 zeolite Si/Al ratio Brønsted/Lewis ratio crystallinity crystal size (µm) particle size (mm) apparent density (g cm-3) BET surface area (m2 g-1) pore volume (cm3 g-1) micropore volume (cm3 g-1)b

catalysta

24 2.9 97% 6.3 0.94 420 0.65 0.17

0.3-0.5 1.21 124 0.43

pore volume distribution of the catalyst (vol %) dp < 10-3 µm 8.1 10-3 e dp e 10-2 µm 14.7 -2 10 e dp e 2 µm 77.2 zeolite acidity measurements NH3 g-1)

total acidity (mmol of base temperature peaks in TPD (K)

0.51 695

tert-butylamine 0.46 577

a Composition: zeolite, 25 wt %; bentonite, 30 wt %; alumina, 45 wt %. b 99% of volume consists of pores with diameter < 0.7 nm.

and follows Mobil patents.39,40 The zeolite was subjected to an agglomeration process with bentonite (Exaloid), using fused alumina (Martinswerk) as an inert charge. The properties of the HZSM-5 zeolite and of the catalyst are presented in Table 1. The Brønsted/Lewis acid site ratio was measured by FTIR spectrometry from the intensity of the adsorption bands at 1550 and 1455 cm-1. Prior to use, the catalyst was calcined at 843 K for 2 h (for the experimental results to be reproducible under reaction-regeneration cycles under isothermal conditions).41

A previous study was carried out to choose the model for simulating the reaction step.37 The best performance is given by the model described below. The mass conservation equation for each lump of the kinetic scheme of the MTG process (Figure 3), assuming plug flow and expressing the concentrations as mass fractions (based on organic components), Xi, is

)

XirW ∂Xi FmT ) ri + ∂t 1 + XW FMoFg mT ∂Xi Xi ∂XW Xi ∂XW (1) 2 1 + XW ∂t 1 + XW ∂z πR F ∂z i

(

g

)

where the equations for the reaction rates, ri, for the different lumps of the scheme in Figure 3 are

rM )

dXM d(W/FMo)

)

aD[-k1XM2 + (k1/Keq)XDXW] - a(k2 + k5XC)XM (2) 1 + kWoXW rD )

dXD d(W/FMo)

rC )

dXC d(W/FMo)

(k2XM + k3XD - k4XC2 ) (k5XM + k6XD + k7XG)XC + k8XG)a 1 + kWoXW

)

aD[k1XM2 - (k1/Keq)XDXW] - a(k3 + k6XC)XD (3) 1 + kWoXW

(4)

For water, the mass conservation equation is

mT ∂XW FmT ∂XW ) rW ∂t FMoFg πR 2F ∂z i

Reactor Simulation

(

Figure 3. Reaction steps in the MTG process.

(5)

g

The term XW is the ratio between the mass flows of water, mW, and of organic components, mO, in the reaction medium. The mass flow of water in the reaction medium is the sum of the mass flow of water formed as a product, mWf, and the mass flow of water in the feed, mWo

mW ) mWf + mWo ) mO

mWf + mWo mO

(6)

On the other hand, the global mass balance must be fulfilled

mT ) mO + mW

(7)

Combining eqs 6 and 7, one obtains

mWf + mWo mT mO mW ) mWf 1+ mO

(8)

In eq 8, the relationship between the mass flows of water formed and of organic components is calculated by taking into account the stoichiometries of methanol dehydration (first step in the scheme of Figure 3) and of hydrocarbon formation [with a general formula (CH2)n] in the lump of oxygenates

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(

)

mWf XD XC + XG ) 18 + mO 46 14

Table 2. Kinetic Parameters2,5

(9)

In eqs 2-4, the deactivation for the dehydration step (first step in the scheme of Figure 3) is quantified by means of an activity, aD, which is different from the activity of the remaining steps of the kinetic scheme in Figure 3, a. These activities are defined as

aD ) a)

(ri)j (rio)j

(ri)1

(ri)1′

) (rio)1 (rio)1′ for j ) 2, ...., 8

Kinetic Parameters for the Main Reactiona 1 (T1 - 673 )] [ 1 1 k ) 11.98 ((2.16) exp[-12 300 ((850)( T 673)] k1 ) 88.73 ((7.89) exp -6400 ((500) 2

k3 ) 9.95 ((3.29)k2

[

k4 ) 9.42 ((3.07) exp -4600 ((1650)

1 (T1 - 673 )]

(10)

[

k5 ) 34.82 ((6.80) exp -10 800 ((750)

(11)

k6 ) 7.62 ((1.60)k5

1 (T1 - 673 )]

1 (T1 - 673 )] 1 1 ) 0.60 ((0.10) exp[-15 350 ((6900)( T 673)]

[

k7 ) 6.893 ((4.90) exp -13 300 ((1950)

In this way, the fact that the deactivation of the methanol dehydration step is significantly slower than the deactivation of the subsequent steps of the kinetic scheme is taken into account. This deactivation is slow because the dehydration occurs even on very weak acid sites and a small number of these sites is sufficient.42 The selective deactivation is a consequence of the different acid strengths of the active sites required in the different individual reactions and has been demonstrated in other reactions such as the skeletal isomerization of n-butenes on chloride alumina catalysts.43 Nevertheless, selective deactivation has not been proposed for the subsequent steps of the kinetic scheme in Figure 3, as it has previously been stated that there is no clear evidence of the intervention of sites of different acidic strengths in the following steps of the kinetic scheme (conversion of oxygenates to light olefins and transformation of olefins into hydrocarbons in the range of gasoline).44 The deactivation kinetic equations are

kdAXA + kdCXC + kdGXG da )a dt 1 + kWXW

(12)

aD ) an where 0 < n < 1

(13)

With these equations, the deactivation for all of the reaction steps is assumed to depend on the same variables (concentrations of components, temperature, and time on stream) and to follow similar kinetics, although aD decreases with time on stream more slowly than does a. Equation 13 relates the two activities in a simplified form. The calculation of the kinetic parameters for the main reaction, eqs 2-4, and for deactivation, eqs 12 and 13, was carried out by fitting the experimental data of composition vs time obtained under different operating conditions [temperature range between 300 and 450 °C, space time up to 0.37 (g of catalyst) h (g of methanol)-1, water/methanol ratios in the feed between 0 and 1] to the values calculated by solving the corresponding mass conservation equations for each lump or component of the kinetic scheme. The procedure is described in detail in previous papers.3,5 In Table 2, the values of the kinetic parameters of eqs 2-4, 12, and 13 are provided. It is noteworthy that the kinetic models for the main reaction, eqs 2-4, and for the deactivation, eqs 12 and 13, take the attenuating effect of water into account, which is quantified by means of the constants kWo and kW, respectively. Judging from the Langmuir-Hinshel-

k8

kWo ) 1.0 ((0.3)

Parameters for the Kinetic Deactivation Modelb 1 (T1 - 673 )] [ 1 1 ) 0.138 ((0.043) exp[-8700 ((1050)( T 673)] 1 1 ) 0.075 ((0.028) exp[-10 650 ((950)( T 673)] 1 1 ) 2.392 ((0.061) exp[-6900 ((450)( T 673)]

kdA ) 1.639 ((0.022) exp -10 800 ((750) kdC kdG kW

n ) 0.25 ((0.01)

Parameters in the Heat Balancec cp ) 0.91 J g-1 K-1 hw ) 11.6 J s-1 m-2 K-1 Ka ) 0.402 J s-1 m-1 K-1 U ) 2.02 J s-1 m-2 K-1 a Steps of the scheme in Figure 3, eqs 2-4. b Equations 12 and 13. c Equations 14 and 15.

wood-Hougen-Watson postulates for ascertaining the role of the active sites in the kinetics of the catalytic reactions, the physical meaning of the parameter kWo is related to the adsorption equilibrium constant of water on the acid sites of the catalyst. On the other hand, in Table 2, it is observed that the constant kW increases as the temperature is increased. This result agrees with the possible physical effect of coke stripping caused by water and also with the hypothesis that both coke formation and the growth mechanism are attenuated by water competing with coke precursors for adsorption on the acid sites. Two energy balances have been considered

For the bed and gas 2

πRi FcpFd mT

∂T

nl

)

∂t nl

(

XirW

∑i ri - 1 + X

)

W

∂T

πRi2FmT

∆Hi

FMo

1 + XW

-

∂T

∑ Xicpi ∂z - mincpin ∂z - 2πRihw(T ) i

(1 + XW

Tw) + KaπRi

2

∂2 T ∂z2

(14)

Ind. Eng. Chem. Res., Vol. 40, No. 26, 2001 6091

For the reactor wall π(Re2 - Ri2)Fwcpw

∂Tw ) 2πRihw(T - Tw) ∂t 2πReU(Tw - Te) (15)

By using these two equations, although they correspond to a homogeneous model, the wall temperature, Tw, and the outside temperature, Te, are differentiated. The heat transfer between the solid and gas and the wall is quantified by the coefficient hw, and that between the wall and the outside by means of the overall coefficient U. The equations for mass and energy conservation are solved with the following boundary conditions for z ) 0

Xi ) Xio

(16)

T ) To

(17)

Tw ) To

(18)

Table 3. Operating Conditions for the Three Test Cases Used for Verifying the Reactor Simulation Program FMo To Q N2 W/FMo case (°C) XWo (g min-1) (cm3 min-1) [gcatalyst h (gmethanol)-1] 1 2 3

350 350 350

0 0 1

1.23 0.48 1.25

138.6 213.6 138.6

0.226 0.573 0.445

The set of partial differential equations (with respect to time and space) comprising the mass conservation equations (eq 1 for n - 1 lumps and eq 5 for water) and the energy conservation equations (eqs 14 and 15) were solved simultaneously with the kinetic equations for deactivation (eqs 12 and 13) by means of a program in Fortran, which provides a discrete longitudinal coordinate and transforms the original partial differential equations into ordinary differential equations. The subroutines DSS032 and DSS044 of the DSS2 library were used to obtain the first and second derivates, respectively, with respect to the longitudinal coordinate. The number, N, and the positions, ξi, of the collocation points required for solving the conservation equations were established taking into account the pronounced concentration profile at the reactor inlet. The collocation points were calculated by means of the following expression

ξi ) ξi-1 + fξ(i-1)∆ξ

(19)

where

∆ξ )

Z fξ - 1 -1 fξ - 1 N

(20)

fξ is a parameter whose value is slightly greater than 1 that sets the spacing between the collocation points. To reach a compromise between the calculation time and the solution accuracy, suitable values for the parameters were determined to be N ) 15 and fξ ) 1.3, which gives rise to the following collocation points: 0.0000, 0.0023, 0.0054, 0.0094, 0.0145, 0.0212, 0.0299, 0.0412, 0.0560, 0.0751, 0.1000, 0.1323, 0.1743, 0.2290, and 0.3000. By solving the set of equations, the longitudinal composition and temperature profiles can be obtained for different values of time on stream. The concentration of lump G, XG, is calculated by difference from unity because the following condition is fulfilled for the organic components in the reaction medium

ΣXi ) 1

(21)

Figure 4. Evolution with time on stream of the temperature profile along the reactor. Dashed lines, experimental results. Solid lines, calculated by solving the simulation model. (a) Case 1 of Table 3. (b) Case 2. (c) Case 3.

In Table 2, the values of the parameters used for solving the model are presented.37 Some of these parameters were determined experimentally (the heat capacity of the bed, cp; the effective axial thermal conductivity of the bed, Ka; and the global heat transfer coefficient, U). The reaction heat of each individual step, ∆Hi, and the heat capacity of each lump, cpi, were calculated from the values of standard components. Results Verification of the Simulation Model. To confirm the adequacy of the model for simulating the laboratory unit, three experimental systems were chosen at atmospheric pressure with the conditions shown in Table 3. These conditions correspond to qualitatively different temperature profiles, and the dilution of methanol in the feed with inert gas or water is taken into account.

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Figure 6. Evolution with time on stream of the activity profile along the reactor. Case 1 of Table 3.

Figure 5. Evolution with time on stream of the product composition at the reactor outlet. Points, experimental results. Solid lines, calculated by solving the simulation model. (a) Case 1 of Table 3. (b) Case 2. (c) Case 3.

The reactor inlet temperature is set at 350 °C, which is approximately the minimum temperature required for a sufficiently fast reaction rate. Furthermore, when this inlet temperature is used, the limiting temperature of 450 °C (to avoid irreversible deactivation by dealumination when a significant water concentration is present)28,29 is not reached at any position in the reactor over a wide range of operating conditions studied. Figure 4 shows the evolution with time on stream of the temperature profile along the reactor. In Figure 5, the evolution with time on stream of the product composition at the reactor outlet is plotted. The dashed lines in Figure 4 and the points in Figure 5 are the experimental results, whereas the solid lines were calculated by solving the simulation model. Adequate fitting between the calculated and the experimental results confirms the validity of the simulation model. It is noteworthy that, in the simulation program used, the past history of the catalyst at each position of the reactor is rigorously taken into account. Several authors have pointed out the importance of this aspect.45,46 As an example of the results of these calculations, Figure 6 shows the evolution with time on stream of the activity profile along the reactor for the experimental system corresponding to case 1 of Table 3.

Figure 7. Evolution of the coke content profile along the reactor. Case 1 of Table 3 (time on stream of 6.4 h).

Figure 8. Evolution with time on stream of the profile along the reactor of water content in the reaction medium. Case 1 of Table 3.

Figure 6 shows a pronounced minimum activity in the lower half of the reactor (at the entrance) from the initiation of the reaction. This minimum is a consequence of the facts that the temperature profile has a maximum at this position in the reactor and the effect of temperature on the deactivation kinetics is more important than the effect of reaction component con-


Figure 9. Evolution with time on stream of the temperature profile along the reactor for two values of space time. XWo ) 0.

Figure 10. Evolution with time on stream of the product composition at the reactor outlet for two values of space time. XWo ) 0.

centration. The maximum temperature is nearer to the inlet for low values of time on stream, which justifies the more severe deactivation of the catalyst at the reactor inlet. The activity profile in Figure 6 is in agreement with the profile of the coke deposited on the catalyst, which is shown in Figure 7. The coke content was determined by combustion in a thermobalance (STD 2960 from TA Instruments) of deactivated catalyst samples taken at different longitudinal positions in the reactor. Prior to coke combustion, the deactivated catalyst sample is subjected to sweeping at 823 K with a He stream at a flow rate of 40 cm3 min-1 for 1 h. Thus, the coke H/C ratio at all longitudinal positions along the reactor is averaged.47 By comparing Figures 6 and 7, it is observed that the minimum activity values correspond to reactor positions in which the coke content in the catalyst is maximum. In the calculation of activity evolution, the kinetic model for deactivation depending on the concentrations of the lumps of the kinetic scheme of the MTG process and on the concentration of water, eq 12, is considered. This water concentration plays an important role in the attenuation of the deactivation and of the kinetics of the reaction steps. Some of the results calculated for the evolution with time on stream of the water composition profile along the reactor are plotted in Figure 8. These results correspond to case 1 of Table 3, in which pure methanol is fed into the reactor.

Study of Process Operating Conditions. Figure 9 shows that, as the space time is increased from 0.114 (g of catalyst) h (g of methanol)-1 (upper graph) to 0.448 (g of catalyst) h (g of methanol)-1 (lower graph), the temperature maximum at each longitudinal position is reached for higher values of time on stream. Furthermore, the maximum temperature values become smaller as the space time is increased. These results are explained by the increase in the total gas flow as the space time is increased, as shown in Figure 10, where the evolution with time on stream of the composition at the reactor outlet is plotted for the same experimental systems of Figure 9. As observed in the upper graph of Figure 10, after 7 h of operation, the formation of the gasoline lump (lump C5+) decreases sharply, and after 11 h of operation, the formation of dimethyl ether also decreases. Nevertheless, for a four times greater space time (lower graph), the conversion of oxygenate compounds is almost complete after 14 h of reaction. A more detailed explanation of the effect of space time is obtained by comparing the results of Figure 11 [for 0.114 (g of catalyst) h (g of methanol)-1] with those of Figure 12 [for 0.440 (g of catalyst) h (g of methanol)-1] in which the evolution with time on stream of the longitudinal profile is plotted for three variables: the water content in the reaction medium (upper graphs), the activity (middle graphs), and the gasoline lump (C5+) content in the reaction medium (lower graphs). These

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Figure 11. Evolution with time on stream of the profiles along the reactor of the water content in the reaction medium (upper graph), the activity (middle graph), and the gasoline lump (C5+) content (lower graph). W/FMo ) 0.114 (g of catalyst) h (g of methanol)-1.

Figure 12. Evolution with time on stream of the profiles along the reactor of the water content in the reaction medium (upper graph), the activity (middle graph), and the gasoline lump (C5+) content (lower graph). W/FMo ) 0.448 (g of catalyst) h (g of methanol)-1.

results correspond to the same experimental conditions as those of Figures 9 and 10, in which no water is fed into the reactor. It is observed that, as the space time is increased, water formation increases, which contributes to the attenuation of catalyst deactivation. Figures 13 and 14 show the effect of the water content in the feed on the evolution with time on stream of the temperature and composition of the product stream, respectively.

In Figure 13, it is observed that the profile of temperature with time on stream is less pronounced and that it becomes flatter as the water content in the feed is increased. Although the temperature decreases as the water content in the feed is increased, the deactivation is lower, and the result is that the evolution with time on stream of the product composition is maintained (lower graph in Figure 14) up to a higher time on stream than for the feed without water (upper graph).


Figure 13. Effect of water content in the feed on the evolution with time on stream of the temperature profile along the reactor, Upper graph, XWo ) 0. Lower graph, XWo ) 1. W/FMo ) 0.223 (g of catalyst) h (g of methanol)-1.

Figure 14. Effect of water content in the feed on the evolution with time on stream of the product composition at the reactor outlet. Upper graph, XWo ) 0. Lower graph, XWo ) 1. W/FMo ) 0.223 (g of catalyst) h (g of methanol)-1.

When methanol is fed with inert gas (nitrogen), the consequence is the attenuation of the temperature and of the temperature peaks, as shown in Figure 15, where the evolution with time on stream of the temperature profile is plotted for two molar flow ratios of inert gas to methanol in the fed, 0.080 (upper graph) and 0.322 (lower graph). Because of this decrease in temperature and because the proportion of inert gas is higher (as shown in Figure 16 for the same conditions as in Figure 15), the conversion to hydrocarbons C5+ is slightly lower, and the deactivation is slightly more pronounced. This latter result is a consequence of the effect of the small increase in the oxygenate concentration in the reaction medium and of the fact that this effect is relatively more important than the inverse effect (decrease in deactivation) of the temperature drop. As mentioned previously, the temperature in the reactor must be controlled to ensure that it is lower than 450 °C. It has been shown28,29 that, at this temperature, for water contents in the reaction medium above 50 wt %, irreversible deactivation of the catalyst by dealumination of the HZSM-5 zeolite is noticeable. This fact restricts the operation of the MTG process5 and of other reactions in which the presence of a high water content in the reaction medium is a key factor, as is the case with the BTG (bioethanol-to-gasoline) process.48 The aforementioned results correspond to a reactor inlet temperature of 350 °C. Under these conditions, a

temperature of 450 °C in the initial section of the reactor is exceeded only when methanol, either pure or diluted with a very low flow of inert gas, is fed into the reactor (upper graph in Figure 15). Nevertheless, the reactor inlet temperature is a key variable because of this effect on the maximum temperature level in the reactor. Inlet temperatures higher than 350 °C have been tested with the aim of determining the operating conditions for which the temperature of 450 °C is reached at any time or position in the reactor. Figure 17 shows the evolution of temperature with time at the reactor longitudinal position for which the maximum temperature is reached. The results correspond to the reactor inlet temperature of 370 °C. In the upper graph of Figure 17, it is observed that the water content in the feed greatly affects the evolution of temperature with time, but the attenuation of the maximum temperature is very small. Consequently, a high water content in the feed (XWo ) 3) is required for the system studied to not exceed 450 °C. On the other hand, the dilution of methanol with an inert gas does efficiently attenuate the maximum temperature in the reactor. As shown in the lower graph of Figure 17, 450 °C is not exceeded for a molar ratio of 0.29 for inert gas to methanol in the feed.

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Figure 15. Effect of inert gas content in the feed on the evolution with time on stream of the temperature profile along the reactor. Upper graph, nin/nMo ) 0.080. Lower graph, nin/nMo ) 0.322. W/FMo ) 0.223 (g of catalyst) h (g of methanol)-1.

Conclusions The good performance of the model proposed in this paper for simulating the laboratory unit is due to the improvements in the kinetic models used to describe the MTG process and the deactivation by coke of the catalyst. On one hand, the consideration of water content in the reaction medium is required because it attenuates the steps of the kinetic scheme of the MTG process. On the other hand, in the kinetic model of deactivation, in addition to the attenuating effect of water in the reaction medium, the dependence on the concentration of the lumps of the kinetic scheme must be taken into account. Once these improvements were incorporated into the kinetic modeling, it was demonstrated that the model proposed in this paper for simulating the reactor is a useful tool for studying the effect of process conditions (space time, dilution of methanol with nitrogen, dilution with water, and feed temperature at the reactor inlet) on the temperature profile along the reactor and on the evolution of the product yields with time on stream. It is noteworthy that, whereas methanol dilution with water in the feed is efficient for lessening deactivation, the dilution of methanol with nitrogen is more efficient for reducing the temperature peaks in the reactor.

Figure 16. Effect of inert gas content in the feed on the evolution with time on stream of the product composition at the reactor outlet. Upper graph, nin/nMo ) 0.080. Lower graph, nin/nMo ) 0.322. W/FMo ) 0.223 (g of catalyst) h (g of methanol)-1.

Acknowledgment This work was carried out with the financial support of the Ministry of Education and Culture of the Spanish Government (Projects CICYT PB96-1478 and CICYT PB97-0644). Notation A, C, D, G, M, W ) oxygenates (methanol and dimethyl ether), light olefins (ethene and propene), dimethyl ether, remaining hydrocarbons, methanol, and water, respectively a, aD ) catalyst activity for the steps of hydrocarbon formation and methanol dehydration, defined as the ratios of reaction rates in eqs 10 and 11, respectively cp ) average heat capacity of the bed, J kg-1 K-1 cpin ) heat capacity of the inert gas, J kg-1 K-1 cpi ) average heat capacity of lump i in the gaseous product stream, J kg-1 K-1 cpw ) heat capacity of the reactor wall, J kg-1 K-1 Fd ) catalyst dilution factor, (kg of bed) (kg of catalyst)-1 FMo ) mass flow rate of methanol in the feed, kg h-1 fξ ) parameter that sets the spacing between collocation points, eq 20 hw ) coefficient of heat transfer between the bed and the reactor wall, J s-1 m-2 K-1 Keq ) equilibrium constant for methanol dehydration Ka ) effective thermal conductivity in the axial direction, J s-1 m-2 K-1

Ind. Eng. Chem. Res., Vol. 40, No. 26, 2001 6097 To ) inlet reactor temperature, K Te, Tw ) temperatures on the outside and at the wall of the reactor, respectively, K t ) time on stream, s U ) global heat transfer coefficient, J s-1 m-2 K-1 W ) catalyst weight, g Xi ) weight fraction of component i based on the organic components XW, XWo ) water/organic component ratios, in mass, in the reaction medium and in the feed, respectively Z ) reactor length, m z ) longitudinal coordinate in the reactor, m Greek Letters ∆Hi ) heat associated with the formation of lump i, J mol-1 ) bed voidage F ) apparent catalyst density, kg m-3 Fg ) gas density, kg m-3 Fw ) density of the reactor wall, kg m-3 ξ ) dimensionless longitudinal coordinate in the reactor

Literature Cited

Figure 17. Effect of water content (upper graph) and inert gas content in the feed (lower graph) on the evolution of temperature with time on stream at the reactor longitudinal position for which the maximum temperature is reached. W/FMo ) 0.152 (g of catalyst) h (g of methanol)-1.

kdA, kdC, kdG ) deactivation kinetic constants for coke formation from the lumps of oxygenates, light olefins, and remaining hydrocarbons, respectively, for the steps of transformation of oxygenates into hydrocarbons, h-1 ki ) kinetic constant of step i in the kinetic scheme of Figure 3 kWo ) parameter that quantifies the resistance to the formation of component i in the corresponding reaction step due to the presence of water in the reaction medium kW ) parameter that quantifies the attenuating effect of water on deactivation min ) mass flowrate of inert gas, kg s-1 mO ) mass flow rate of organic components, kg s-1 mT ) mass flow rate excluding the inert gas, kg s-1 mW ) mass flow rate of water in the reaction medium, kg s-1 mWf ) mass flow rate of water formed as a product, kg s-1 mWo ) mass flow rate of water in the feed, kg s-1 N ) number of collocation points Re, Ri ) external and internal radii of the reactor, respectively, m ri, rio ) formation rates of lump i at a given time on stream and for the fresh catalyst, respectively, g h-1 (g of catalyst)-1 (ri)j, (rio)j ) reaction rates for formation of component i in step j of the kinetic scheme, at a given time on stream and for the fresh catalyst, respecively, g h-1 (g of catalyst)-1 T ) temperature, K

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Received for review February 28, 2001 Revised manuscript received September 26, 2001 Accepted October 3, 2001 IE0101893

MTG Process in a Fixed-Bed Reactor. Operation and Simulation of a

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