WO1994000411A1

WO1994000411A1 - Process and apparatus for producing tertiary ethers

Info

Publication number: WO1994000411A1
Application number: PCT/FI1993/000266
Authority: WO
Inventors: Isto Eilos; Juhani Aittamaa; Juha Jakkula
Original assignee: Neste Oy
Priority date: 1992-06-22
Filing date: 1993-06-22
Publication date: 1994-01-06
Also published as: AU4329493A; FI922909A0; FI93207C; FI93207B; FI922909A

Abstract

The present invention relates to a process and an apparatus for producing tertiary alkyl ethers from isoolefins and alcohols. According to the process, the isoolefins and alcohols are reacted with each other in a reaction space (31) in the presence of a cation exchange resin for forming the ethers. The process is carried out in an reactor vessel comprising a lower section in which a fluidized-bed section (35) can be formed, and an upper section in which a fixed-bed section (36) can be formed, respectively. The inlet nozzle (30A) for initial reactant mixture feed is adapted to the reactor vessel lower section and the outlet nozzle (30B) for removal of reaction products is adapted to the reactor vessel upper section, respectively. According to the invention the fixed reaction bed (36) is advantageously achieved by using a reactor equipped with an expansion section (37) above which the reactor cross-sectional area is increased sufficiently to reduce the superficial velocity of the liquid upflow from the reactor lower section to the upper section below the minimum fluidization velocity. By virtue of the invention, the catalyst contained in the fixed bed can be changed even during run by temporarily reducing the superficial liquid velocity so much as to allow the catalyst particles contained in the fixed bed to fall by gravity to the lower section of the reactor, wherefrom the spent catalyst can be unloaded as necessary and replaced by new catalyst.

Description

Process and apparatus for producing tertiary ethers

The present invention relates to a process according to the preamble of claim 1 for producing tertiary alkyl ethers from isoolefins and alcohols.

According to the process, C₄ to C₇ isoolefins and even higher isoolefins, as well as mixtures thereof, are reacted with lower aliphatic alcohols, in particular methanol or ethanol, to produce corresponding ether compounds. The invention also relates to an apparatus according to the preamble of claim 9 for producing tertiary ethers and to a method according to the preamble of claim 24 for operating such an apparatus. The apparatus according to the invention comprises a contiguous and elongated reactor vessel filled with cation exchange resin, the longitudinal axis of said reactor vessel being essentially vertically aligned and said reactor vessel having an inlet nozzle for feed of initial reactants and an outlet nozzle for removal of reaction products.

With regard to ternary ethers, the present invention in particular relates to the production of methyl tert-butyl and ethyl tert-butyl ethers (hereinafter called MTBE and ETBE, respectively) as well as tert-amyl methyl and tert-amyl ethyl ethers (hereinafter called as TAME and TAEE, respectively).

The etherification reaction of isoolefin with alcohol is exothermic (heat releasing) by nature. Thus, the choice of the reactor type with mass production in mind can be any of the following: fixed-bed reactor, tubular reactor, fluidized-bed reactor or reactive distillation reactor. The latter has lately been favoured to an increasing breadth. This is because distillation provides reaction of the reactants and separation of reaction products therefrom simultaneously. Such a property is advantageous in terms of reaction kinetics as the process involves an equilibrium reaction in which the maximum conversion is determined by the thermodynamical equilibrium of the reaction system. So, when MTBE for instance is produced by the conventional way of combining a fixed-bed reactor with distillation separation of the product, typically approx. 90 to 95 % conversion of isobutene is achieved while the reactive distillation reactor provides a conversion in the range from 98 to 99 % . Conventionally, the process catalyst employed is a sulfonated polystyrene/divinylbenzene-based cation exchange resin having a particle size of 0.1 to 1 mm typical. Owing to channelling of the flows, a particle of this size exhibits unsatisfactory hydrodynamics in a large, industrial-scale distillation unit, which necessitates the use of a larger catalyst charge. However, a large catalyst charge only worsens the above-mentioned channelling problem. Therefore different types of bag/sock systems have been developed, and attempts have been made to shape the catalyst particles identical to the packing pieces. Of the latter approach, the only industrially implemented application is realized by NPO Yarsintez and is disclosed in U.S. Patent Specification No. 3,965,039.

Another significant problem of catalytic distillation is associated with catalyst charge change. Because the catalyst is located within the distillation column, changing catalyst charge is a significantly more complicated operation than in a conventional fixed-bed reactor and cannot be performed on-line during run. This problem can be alleviated through the use of a prereactor that acts as a guard by absorbing the catalyst poisons and is provided with an easier catalyst charge change facility.

As concerns the etherification reaction, it is important to control the alcohol/ isoolefin ratio, since during operation with an alcohol-lean mixture the isoolefin is dimerized and/or polymerized in the catalyst bed, and with an alcohol-rich mixture, two alcohol molecules mutually form a corresponding ether. In a reactor preceding the distillation phase, use of excess alcohol causes problems in the product separation by distillation as the alcohol forms an azeotrope with both the product and the residual hydrocarbons. Therefore, the prereactor of the reactive distillation stage is conventionally operated with an alcohol-lean mixture. In a conventional process having typically two prereactors in series, the process is usually run with a slightly alcohol-rich mixture.

The hydrocarbon stream is fed into the prereactor maximally concentrated with respect to the isoolefm. As the hydrocarbons contained in the feedstock have been prevented up to this step from reacting into ether in any preceding stage, the formation of heat is maximal particularly in this process phase. Consequently, the prereactor employed in the process is preferably a fixed-bed reactor provided with a recycling facility, or alternatively, a tubular reactor provided with cooling. The prereactor can also be a fluidized-bed reactor. Such an embodiment is described in the DE Published Patent Application No. 2,944,914.

In a fluidized-bed reactor the catalyst is fluidized by the flow of the hydrocarbon feedstock in liquid phase. The flow rate required for fluidization is maintained by recycling. If the feedstock flows contain catalyst poisons in significant amounts, the fluidized-bed reactor can be utilized for removal of such adverse compounds before the process flow enters the reactive distillation column. The use of a fluidized-bed reactor also offers the important benefit that the catalyst charge can be changed continuously during run. Moreover, control of temperature gradients in a fluidized-bed reactor is more readily possible than in any other reactor type.

In the embodiment described in the DE Published Patent Application No.

2,944,914 the reaction mixture is routed from the fluidized-bed reactor to another reactor in which the catalyst is contained in a fixed bed. This prior-art embodiment suffers from certain essential disadvantages. For instance, the catalyst of a fixed-bed reactor cannot be changed without resorting to process shutdown in any other manner except by arranging multiple parallel reactors subsequent to the fluidized-bed reactor, which arrangement in turn increases the purchasing and operating costs of the process apparatus. Moreover, catalyst particles escaping the fluidized-bed reactor easily clog the piping between the fixed-bed reactor and the fluidized-bed reactor. Loss of catalyst particles by through-flow occurs particularly in disturbance situations in which the flow velocity changes abruptly, as well as in situations when the fluidized bed is fed with an excess amount of the catalyst or the catalyst particle size is smaller than the size used as the design basis for the reactor. It is also known in the art to combine several reactor beds within one reactor body. Thus, U.S. Patent Specification No. 4,589,927 anticipates a liquid fluidized reactor containing two fluidized beds, namely a first dense bed in the bottom of the reactor body containing coarse solid particles and a second bed in the upper section of the reactor body containing fine solid particles. In order to improve the separation of the small and large particles, the upper section of the reactor body is enlarged and joined to the lower section via a collar section. The fine solid particles are recirculated from the upper region through the reactor, a particle separator, an external regenerator and conduits back through the dense bed of large particles.

The use of the reactor described in the U.S. Patent No. 4,589,927 for continuous chemical and biological reactions, such as waste water treatment, starch hydrolysis amino acid production, and production of enzymes or gene products, is suggested. No reference to etherification is made, and it is obvious that no significant improvement of the conversion rate of the etherification reaction can be obtained with a reactor containing two fluidized beds. Furthermore, any etherification product produced with this prior art reactor will contain entrained fine particles which will have to be carefully removed by an additional separation step.

It is an object of the present invention to overcome the disadvantages associated with the prior-art technology and to achieve a novel process and apparatus for producing tertiary ethers. The invention is based on the concept of performing the reaction of isoolefins with alcohol at least partially in an essentially vertical, contiguous reaction vessel having a lower section suited for formation of fluidized bed and an upper section suited for formation of a fixed reaction bed. According to the invention the fixed reaction bed is here obtained by limiting the superficial velocity of the liquid flowing through the reactor so that said velocity at the upper section of the reactor remains below the minimum fluidization velocity. By arranging the fixed bed above the fluidized bed, the change of the catalyst contained in the fixed bed can be carried out by temporarily decreasing the flow velocity sufficiently to allow the catalyst particles to fall by gravity from the fixed bed to the lower section of the reactor, wherefrom the spent catalyst is unloaded as necessary and replaced by new catalyst. Conversely, the catalyst particles falling to the lower section of the reactor can be left there to form the fluidized bed.

More specifically, the process according to the invention is mainly

characterized by what is stated in the characterizing part of claim 1. In the reactor according to U.S. Patent No. 4,589,927, the whole solid particle bed is fluidized, which means that the flow velocity at each point of the reactor body is higher than the minimum fluidization velocity. In contrast, in the reactor according to the present invention, the flow velocity in the upper section is smaller than the minimum fluidization velocity of the particle mixture contained therein. In addition, the volume of the upper section is in the range from about 20 to about 95 % of the total volume of the reactor. Typically the upper section makes up 40 to 80 %, preferably about half or more of the entire reactor volume. This reactor design means that all of the catalyst particles cannot be contained in the lower section, and it will make it easier to form a fixed bed in the upper part of the reactor. When the flow velocity is zero in a reactor according to the invention, the entire bulk of particles will form a fixed bed and at least some of the particles will be in the upper section of the reactor. When the velocity in the lower section is increased above the miniumum fluidization velocity of the catalyst particles, the catalyst bed in the lower section will start to expand and a part of the particles spreads to the upper section. Since the flow velocity in the upper section is below the minimum fluidization velocity of the mixture of particles present in the upper section, no segregation of the smaller particles will take place. In this context it should be noted that the minimum fluidization velocity depends on the avarage particle size. According to the invention, the minimum fluidization velocity will remain higher than the linear flow velocity of the liquid flowing through during the process.

The lower section is preferably connected to the upper section by a conical expansion section. To avoid the formation of a spouted bed in the upper section, the cone angle is 1 ° to 45°, preferably from about 5° to about 30°. The outlet nozzle is fitted in the upper part of the reactor above the fixed bed. Therefore, the product removed from the reactor will contain only small amounts of catalyst particles, if any.

The apparatus according to the invention for producing tertiary ethers is characterized by what is stated in the characterizing part of claim 9 and the method according to the invention for operating the claimed apparatus is characterized by what is stated in the characterizing part of claim 24.

In the context of the present invention, any reactor preceding the actual product separating distillation stage in the process is called a prereactor whether or not the product separation by distillation following the preceding reactor takes place by reactive distillation. In certain cases, particularly in the production of TAME and TAEE, the product obtained from such a prereactor need not be separated by distillation from the hydrocarbon stream, since both the product and the hydro-carbon stream are directly suited for use as end product components (namely, as TAME and TAEE are concerned, said reaction products and the C₅ hydrocarbon stream are suitable as motor gasoline components).

Reactants used in the production of tertiary ethers are isoolefins containing 4 to 7 carbon atoms, such as isobutene, isopentene, isohexene and isoheptene or hydrocarbon streams containing mixtures thereof, combined with lower alcohols (that is. alcohols containing 6 carbon atoms or less) such as methanol, ethanol, propanol, etc. The hydrocarbon feedstock also frequently contains n-olefins and saturated hydrocarbons.

The term "particle" (such as an ion-exchange resin or catalyst particle) in the context of the present patent application refers to particles having a suitable bead size for liquid fluidization. The acid cation exchange resin employed as catalyst according to the invention has a bead size in the range from 0.01 to 10 mm, preferably approx. from 0.1 to 1 mm. The term reaction bed refers to a space in which the catalyst and the liquid are contacted and so terms "reaction bed" and "reaction layer" are used

synonymously within the context of this application.

The term "liquid flow velocity" or the superficial liquid velocity (average velocity) is defined as the volume flow rate of the liquid divided by the cross section of the empty column.

Discussions related to reaction zones hereinafter refer to such expressions as "at least one single reaction zone" and "at least one second reaction zone", which must be understood to indicate that according to the invention the reactor can be provided with a plurality of subsequent liquid-fluidized and fixed beds, respectively. The fact that a fluidized bed can be divided into subsections with the help of flow redistributors is known in the art for example from the textbook Kunii, D. and Levenspiel, O., Fluidization Engineering, 2nd Ed. (1977), Robert E. Krieger Publishing Co. New York, pp. 26 - 34.

This invention optimizes the flow conditions in MTBE/ETBE/TAME/TAEE processes for the first reactor preceding the process stage of product separation by distillation so that the catalyst in said reactor is present in both the fluidized bed and fixed bed phases. Then, the reactor bottom is preferably provided with a catalyst bed exhibiting a high expansion factor (over 50 %) and a high degree of mixing, while the catalyst bed above this bed has the properties of a conventional liquid-fluidized bed. namely a normal expansion factor (20 to 30 %) and a moderate degree of mixing. The upper section of the reactor is provided with a fixed catalyst bed and, in order to achieve high conversion rates, the volume of the upper section is large in comparison to the lower section..

All these layers can in principle be implemented in said reactor either as single or separate layers. The essential requirement according to the invention is, however, that the reactor construction must be capable of supporting a normal liquid-fluidized bed, and above this, a reaction bed functioning at least essentially as a fixed bed.

According to the invention the lower section and the upper section of the reactor are preferably filled with the same type of catalyst, typically the particles of the above-described cation exchange resin such as is used in a conventional manner in state-of-the-art methods for producing tertiary ethers.

As mentioned above, the lower section is joined to the upper section via an expanding section, the cross-sectional area of the reactor vessel increasing above the expanding section so much as that the superficial velocity of the liquid entering the upper section from the lower section falls below the minimum fluidization velocity. The expanding section comprises for instance a collar section shaped as a truncated cone with a cone angle of 1° to 45°, preferably 5° to 30°. Then, the cross-sectional area of the reactor vessel at the upper section is approx. 2- to 10-fold the reactor vessel cross-sectional area at the lower section.

The upper section in this embodiment is so spacious that it incorporates at least a portion of the expanding section.

It must be noted herein that the art conventionally employs gas-fluidized-bed reactors which have an expanded upper section, In these, however, the reactor upper section performs a function different from that in the present invention, Namely, in the art the expanding section has not been designed for forming and maintaining a fixed bed, but rather, its function is to minimize the carry-over of entrained particles from the reactor. In a liquid-fluidized bed, the carry-over of particles from the bed becomes no problem as the fluidization by its nature does not take place via intense bubble formation, but instead, the bed upper surface is entirely smooth. Then, it is possible to utilize the expanding section effectively as the expansion space of the bed.

According to another embodiment of the invention, the superficial velocity of the liquid entering the upper section is reduced below the minimum fluidization velocity by derouting at least a fraction of the liquid flow leaving the fluidized bed into an external recycling loop. Therefore, an inlet nozzle of the recycling line is adapted above the lower section of the reactor vessel, said inlet nozzle permitting recycling of at least a fraction of the product mixture passing through the lower section back to the inlet nozzle of the initial reactants. In addition to accomplishing the required fluidization, the recycling of the liquid flow through the liquid-fluidized bed simultaneously improves conversion in comparison to a once-through fixed bed having heat gradient problems, thus justifying the use of recycling also in the embodiment described above. When applied to the production of tertiary ethers, in both embodiments the superficial liquid mixture velocity is reduced after the fluidized bed to a value which is approx. from 10 to 40 %, preferably approx. 30 to 35 %, of the superficial velocity in the fluidized bed. Depending on the process pressure, the reactor can have a circular or essentially square cross section. This design rule applied to both above-described embodiments. At low process pressures a square or circular cross section offers both technical and economical benefits, while a reactor vessel with circular cross section is preferable at high process pressures. In a preferred embodiment of the invention the lower and upper sections are separated from each other by means of a flow distributor that performs homo-genization of the liquid flow over the entire cross section of the reactor. As mentioned above, an esterification reaction producing tertiary alkylethers such as MTBE, ETBE, TAME or TAEE is exothermic, whereby temperature increase reduces the equilibrium constant of the reaction, and thereby, the theoretical maximum conversion of the reaction. Close to the reaction equilibrium, the presence of the product slows down the reaction rate. Consequently, the reaction proceeds less vigorously than in the situation when the product is absent, that is, in the initial phase of the reaction. To be able to control the temperature gradient, a high bed degree of mixing in this stage is desirable. In a fluidized-bed reactor this is accomplished by way of a high superficial velocity of the fluidizing liquid. Increasing the superficial liquid velocity in the entire reactor to a value required by a high degree of mixing, however, necessitates a high- volume reactor as the bed expands strongly. Moreover, efficient mixing is not desirable in the upper section of the bed as the attainable conversion is reduced thereby. The problem can be overcome by using a conical reactor bottom. This arrangement achieves efficient mixing in the reactor bottom section and eliminates the stagnation of large particles on the flow distributor plate. Suitable cone angle of the conical expanding section is hereby 1 ° to 60°, preferably approx. 1° to 45°, in particular about 1 ° to 30°.

According to a preferred embodiment the reaction mixture leaving the fluidized bed is cooled prior to its entry into the fixed bed. To this end the fluidized bed of the reactor lower section, or alternatively, the space above this is provided with a heat exchanger, for instance a tubular heat exchanger, which simultaneously can perform a flow distributor. The bottom section of the reactor vessel is provided with catalyst inlet and outlet nozzles. Through these nozzles the ion exchange resin can be both fed into the reactor vessel and unloaded therefrom, respectively. According to the invention, a method is provided for operating such an apparatus, in which method the reactor vessel is first loaded with a sufficient charge of catalyst particles to form two reaction beds. Next, the reactor vessel lower section is filled via the reactor vessel inlet nozzle with a liquid mixture of the initial reactants to be reacted, while simultaneously adjusting the liquid flow rate with the help of a flow control element, which is installed to the pipeline leading to the inlet nozzle, to such a high level that the catalyst particles contained in the reactor vessel lower section become entrained thus forming a fluidized bed. With the help of said flow control element, the superficial liquid velocity is further set to such a high level that excess fraction of the catalyst is carried over with the liquid mixture flow to the reactor vessel upper section. The superficial liquid mixture velocity in the reactor vessel upper section is arranged to fall below the minimum fluidization velocity, whereby a fixed bed in formed therein. When necessary during the reaction, a portion of the cation exchange resin contained in the fluidized bed is removed and replaced by new or regenerated resin.

In small-scale test performed using a conically shaped reactor vessel, the catalyst was not found to circulate between the fluidized-phase and fixed-phase beds. This offers the benefit that most of the catalyst poisoning can occur only in the fluidized bed contained in the reactor lower section. Therefore, only this fraction of the catalyst charge need be entirely changed during regular catalyst change. The fraction of the catalyst contained in the fixed bed can be transferred to the fluidized bed zone during the catalyst change, or alternatively, in conjunction with process run. Catalyst change occurs via a catalyst unload nozzle adapted to the lower section of the reactor vessel that can be opened during catalyst change. By using a smaller volume of the catalyst unload compartment, the amount of unloaded catalyst can be limited to the volume of said compartment, the rest of the catalyst remaining in the reactor. After the desired amount of catalyst is removed from the fluidized-bed section, the liquid flow in the reactor is stopped for 1 to 30 min (typically about 5 minutes), whereby the catalyst contained in the reactor upper section falls down by gravity. Next, the liquid flow rate in the reactor is slowly increased by recycling of the liquid above the minimum fluidization velocity until the desired amount of catalyst becomes entrained thus expanding into the reactor upper section to form a fixed bed there. Subsequently, a portion of new catalyst is fed into the reactor and the liquid flow rate is reduced to the desired fluidization velocity.

The invention also provides a process for producing tertiary alkyl ethers from an isoolefin or mixtures thereof and from at least one aliphatic alcohol. According to the method the materials participating in the reaction are processed into a liquid mixture which is fed into the above-described reaction apparatus where the olefins and the alcohol are reacted.

The invention offers significant benefits.

By combining the fixed-bed and fluidized-bed zones in the above-described manner within a single reactor, a vastly simplified catalyst change is provided. Furthermore, the expanded reactor top performs as an effective expansion space for the fluidized bed as during an abrupt bed expansion in a disturbance situation for instance, this zone provides rapid reduction of superficial liquid velocity, whereby the fluidized-bed voidage is reduced. Thus, the total volume occupied by the catalyst bed is reduced.

In summary, the benefits offered by the invention over a conventional fluidized-bed reactor are listed:

- the catalyst of the fixed bed can be changed during run,

- good mixing is attained immediately above the flow distributor,

- no dead settling regions occur in the corners of the flow distributor, because a reactor lower section of the present type in a fluidized-bed reactor achieves recirculation of particles that is directed downward along the walls and upward in the center,

- good material and heat transfer is attained immediately above the flow

distributor in the reactor,

- good mixing is attained in that section of the reactor where desirable. - removal of large particles from resting on the flow distributor plate is attained,

- a high degree of mixing is obtained at reactor bottom without any significant increase in reactor size,

- use of a small bed expansion factor is possible in the reactor upper section, whereby only moderate mixing occurs, thus resulting in improved conversion in the reactor,

- a low degree of mixing is attained in the reactor upper section,

- reaction conversion improvement by several percent is attained,

- effective expansion space is provided in the reactor upper section,

- risk of catalyst carry-over from the reactor is reduced,

- greater tolerance is obtained for the amount of catalyst charge used,

- the amount of spent catalyst to be replaced at a time is reduced, since the particles contained in the fixed-bed and the fluidized-bed sections, respectively, are not mixed with each other during normal run,

In comparison to the reactor described in the U.S. Patent 4,589,927, should that one be used for etherification, the present reactor provides higher conversion rates and a product containing less entrained catalyst particles (usually there is only a minute amount of catalyst particles in the product, if at all).

If the reactor is divided in sections with the help of a heat exchanger, the following extra benefits are gained:

- the reactor lower section can be run at a higher temperature than the reactor upper section,

- higher conversion is attained without increasing the catalyst charge,

- catalyst charge fractions divided by the reactor heat exchanger into a lower zone and an upper zone do not mix with each other,

- deactivation of catalyst principally occurs in the reactor lower section, - amount of catalyst charge fraction subjected to poisoning is reduced, and simultaneously, the amount of spent catalyst to be replaced at a time is reduced. In the following the invention with its details and benefits are explained closer with the help of a detailed description by making reference to the annexed drawing in which,

Figure 1 shows the diagrammatically the truncated cone to which reference is made in conjunction with the explanation of the theoretical background for the invention,

Figure 2 shows diagrammatically a prior-art fluidized-bed reactor,

Figure 3 shows diagrammatically a first embodiment of a fluidized/fixed-bed reactor according to the invention,

Figure 4 shows diagrammatically the above-mentioned first embodiment provided with a sidestream drawoff and associated recycling circuit,

Figure 5 shows a reactor similar to that illustrated in Fig. 3 complemented with a heat exchanger,

Figure 6 shows a reactor similar to that illustrated in Fig. 4 complemented with a heat exchanger,

Figure 7 shows diagrammatically a second embodiment of a fluidized/fixed-bed reactor according to the invention, and

Figure 8 shows the process schematic for a combination catalyst charge load and unload arrangement adapted to the embodiment illustrated in Fig. 3.

First, the theoretical background of the invention will be examined:

Liquid fluidization

Liquid flow upward through a solid particles layer (that is, a fixed bed) can take place at a low superficial velocity without moving grain particles if the density of the particles is greater than the liquid density as is the case in the examples to be described below. When the superficial velocity is increased, the particles commence a motion in confined areas and the bed layer volume will be expanded. At a sufficiently high flow rate the particles will be entrained in the liquid flow. In this phase the frictional force acting between the particle and the liquid flow, that is, the fluidization flow, overcomes the particle weight, thus cancelling the vertical compressive force between superimposed particles, and the pressure drop over the layer becomes equal to the effective weight of the particles per unit area. The superficial liquid velocity at the start of the above-described phase is called the minimum fluidization velocity. When the flow rate is increased in the liquid-fluidized bed so as to exceed the minimum fluidization velocity, the fluidized bed will be expanded homogeneously and the particle concentrations in different points of the bed will be equalized. Such a fluidized phase is called a smooth, homogeneous, or simply, liquid-fluidized phase.

An essential dissimilarity between gas-fluidized and liquid-fluidized beds is the large volume difference between the fluid bubbles (that is, gas bubbles and liquid bubbles). In gas fluidization a significant fraction of the fluidizing gas tends to rise through the fluidized bed in the form of large-volume bubbles with particles following in its wake; simultaneously, the flow is channelled. Owing to the large volume of the bubbles, the upper surface of the gas-fluidized bed is very unstable, and entrained particles will also be found above the bed upper surface. Such a fluidized phase is denoted by its character as aggregative, heterogeneous, bubbling, or gas-fluidized phase. In liquid fluidization the bubbles remain very small, the fluidized bed upper surface is relatively smooth, and no significant amount of entrained particles will be found above the fluidized bed upper surface provided that the superficial liquid velocity is maintained below the free-fall settling velocity of the particles. Liquid fluidization conventionally uses a flow rate capable of achieving 5 to 50 % expansion of the bed. Typically the expansion is in the order of approx. 20 to 30 % . Such a low bed expansion fails to attain complete mixing of the bed solids in reactors with a large height/diameter ratio as is the case in the examples to be described below. The bed exhibits a distinct temperature gradient, and the larger particles will be segregated on the bottom so that the large particles are found at a higher probability close to the flow distributor on the reactor bottom than on the upper surface of the fluidized bed. Owing to turbulence occurring in the bed. the reactor must be considered a mixing reactor in which any reasonable catalyst charge fails to achieve as high conversion as is possible in a fixed bed reactor. Typically, a fixed bed reactor used for production of MTBE, for instance, can achieve a 80 to 90 % isobutene conversion with a reasonable size of catalyst charge, while a fluidized-bed reactor operated in similar conditions achieves 70 to 80 % conversion only.

Semifluidized bed technique A combination fluidized/fixed bed is called a semifluidized bed. In accordance with conventional techniques, the upper section of a liquid-fluidized bed can be provided with a fixed bed by equipping the upper section of the fluidized bed with a particle disengaging separator screen that permits the flow of the liquid alone, while the entrained particles are packed against it. In the present invention the improvement over prior-art techniques is therein that the fixed bed is achieved by reducing the superficial velocity of the liquid exiting the fluidized-bed section to a value below the minimum fluidization velocity. According to a preferred embodiment of the invention, this end is attained by expanding the reactor vessel cross sectional area at the upper section of the bed.

As a background, such a case of conventional technique is first examined in which the minimum semifluidized phase is attained by adjusting the liquid flow rate to a value that makes the expansion of the bed through liquid fluidization to barely elevate the bed height equal to the location level of the separator screen for entrained particles. From this point onward, the liquid flow rate can be increased until the upflow transfer of all particles results in total holdup of solids against the particle separator screen, whereby the bed obviously acts as a distinct fixed bed. The relative fractions of the fixed-bed and the fluidized-bed sections in the system can be estimated by expressing the system material balance as: m_tot = m_pa + m_fl ( 1 ) where m_tot = total mass of particles in the bed

m_pa = total mass of particles in the fixed-bed section

m_fl = total mass of particles in the fluidized-bed section The equation can be rewritten:

(m/h)_ist · h_ist = (m/h)_pa · h_pa + (m/h)_fl · h_fl (2) where (m/h) = mass of particles per unit bed height

h_ist = height of initial fixed bed for zero liquid flow rate

h_pa = height of fixed-bed section formed by upward build-up against particle separator screen

h_fl = height of fluidized-bed section On the other hand

(m/h)_ist = (1-∈_ist)·A·p_p ( 3 ) where ∈_ist = voidage in initial fixed bed

A = fixed bed cross-sectional area

ρ_p = density of a single particle

(m/h)_fl = (1 + (BE/100 %))^-1 · (m/h)_ist (4) where BE = fluidized bed expansion factor from initial state [%] Assuming that

(m/h)_pa = z · (m/h)_ist (5) where z = fixed bed packing factor Combining equations 1...4 gives: hp_a = (h_ist · (1 + (BE/100 %)) - h_r)/(z · (1 + (BE/100 %)) - 1 ) ( 6 ) where h_r = location height of particle separator screen, i.e., total height of semifluidized bed

With the help of Eq. 6, it is possible to calculate the fractions of catalyst contained in the fixed bed and in the fluidized bed, respectively.

In prior-art systems the pressure drop over the fixed bed becomes high, which is a significant drawback.

According to the invention, a semifluidized bed is achieved by reducing the superficial liquid velocity below the minimum fluidization velocity. This end is advantageously attained by making fluidization occur in an upward expanding conical reactor in which the superficial liquid velocity is reduced from the bottom upward. As the superficial liquid velocity changes as a function of the fixed-bed height, the voidage of the fluidized-bed section is not constant.

With reference to Fig. 1, fluidization is next examined in a truncated cone with circular cross section, bottom radius r₀ and the walls expanding upward so as to form an angle α with vertical. The mass of particles per unit height for an entirely fluidized bed can be calculated starting from the equation

where BE₀ = fluidized bed expansion factor for height level zero [%]

(m/V)₀ = particle mass concentration for height level zero A = fluidized bed cross section at height level h and substituting (m/V)₀ = (1 -∈_ist) · p_p/(1 + (BE/100 %)) ( 8 )

Total mass of particles contained in the fluidized bed is calculated from equation:

In the right-side term of Eq. 9 the cross-sectional area and the expansion factor are functions of height level within the fluidized bed. The cross-sectional area is related to the height level by:

A = π · (r₀ + h · tan α)² ( 10 )

For most cases of liquid fluidization at low fluidization efficiency, the expansion of the fluidized bed as a function of superficial liquid velocity can be approximated by a linear function:

(BE/100 %) = a + b · v = a + b · (V/A) ( 11 ) where a, b = constants

v = superficial liquid velocity

V = liquid volume flow rate

Substituting Eqs. 10 and 11 in Eq. 9 and moving the constants in front of the integral, one obtains

Solving Eq.11 for a > -1 gives

As Eq.11 is valid only for a bed in an entirely fluidized phase, Eq.12 is valid for the following condition:

where v_mf = superficial liquid velocity in the minimum fluidization

condition

Assuming the voidage of the fixed bed to be independent on the superficial liquid velocity and equal to the voidage of the fixed bed in the initial state, one obtains

where (m/V)_ist = (1 -∈_ist)r_p (17) For the above-discussed conical reactor section, one obtains m_pa = p_p(1 1-∈_ist)π(p₀ ²(h₂ -h₁) + ρ₀(h₂ ² -h₁ ²)tanα + 3^-1(h₂ ³ -h₁ ³)tan²α) ( 18 )

This equation makes it possible to calculate the total mass of particles contained in the fixed bed when the bed height (h_ist) in the initial state is known for zero superficial liquid velocity. Then, m_pa = m_tot and h₁ = 0. Substituting the formula thus obtained in Eq. 13, the fluidized bed height can be iteratively computed from the equation:

If the fluidized bed height becomes so large that the superficial liquid velocity in the upper section of the bed falls below the minimum fluidization velocity, the bed upper section behaves as a fixed bed. Assuming that bed expansion prior reaching the minimum fluidized state can be approximated by the linear Eq. 11 (same equation as used for approximating bed expansion in the fluidized state), the height of the bed section can be solved by replacing the term h of Eq. 19 by the height obtained from Eq. 14 (i.e., the term v_mf in Eq. 14 is replaced by the superficial liquid velocity at which the bed starts to expand) and then

complementing it with the term representing the fixed-bed section, which can be obtained from Eq. 18 through substituting in the equation the terms h₂ = total height of particle bed and h₁ = height obtained from Eq. 14.

If the height of the fluidized section is limited by means of a separator screen in the conical expanding section of the reactor and the same assumptions on imtial conditions are made as those for deriving Eqs. 6 and 19, then the height of the bed section packed against the particle separator screen can be solved with the help of Eqs. 1, 13 and 18.

In the following, the alternative process embodiments illustrated in Figs. 2 to 8 will be discussed. Figure 2 shows diagrammatically a conventional reactor tube 21 operated in the fluidized-bed state. The reactor is typically employed as prereactor in ether production. Here, the alcohol and olefin components are first combined into a mixture which is fed with the help of a reactor feed pump 22 into the reactor 21. When necessary the reaction mixture is heated by means of a heat exchanger 23. The bed expansion factor can be, for instance, in the range from 5 to 50 % and a suitable superficial liquid velocity is set by adjusting the output of the pump 22. A portion of the reactor exit flow is routed to a catalytic product-separating distillation process, for instance, while the remaining portion is recycled to the inlet nozzle of the reactants in order to improve conversion. The recycling line incorporates a heat exchanger 24 for the cooling of the recycled portion of the product.

Figure 3 shows diagrammatically a first embodiment of the invention. Except for the reactor, the process layout of this embodiment is implemented using the above-described conventional techniques. Accordingly, in this embodiment the liquid reactant mixture is fed via a pump 32 and an optional heat exchanger 33 and an inlet nozzle 30A to the reactor 31, wherefrom the product mixture is removed via an outlet nozzle 30B. A portion of the product mixture is recycled via a heat exchanger 34 back to the inlet nozzle. Principally, the reactor is a reactor tube similar to that employed in the above-described example in having a fluidized bed 35, while, however, its top is provided with a conical expanding section 37, and above this, a second reaction section 36 with a cross-sectional area larger than that of the reactor center section, said second reaction section being capable of forming a fixed bed. The cone angle of the conical expanding section is typically approx. from 5° to 30°, while in principle it can vary in the range 1 ° to 45°. The expansion factor of the fluidized bed is 5 to 50 %, typically approx. 20 to 30 %, and it is set by adjusting the output of the pump 32.

Close to the inlet nozzle of initial reactants, the reactor bottom section is provided with a flow distributor 38 by means of which the liquid flow is divided evenly over the reactor cross section. The flow distributor can be, for instance, a perforated plate dimensioned according to normal fluidized bed fluid distributor design rules. The sizes of holes in the flow distributor must be such that permit downward flow of catalyst particles at shutdown of reactor flow. The bottom of the reactor 31 is shaped into a conical expanding section 39. This shaped form of the reactor bottom achieves good mixing and eliminates settling of large particles on the flow distributor plate. The cone angle of the conical expanding section 38 is here 5° to 60°, preferably approx. 5° to 30°.

The flow rate required in the reactor is determined by the maximum particle size of the catalyst and the expansion factor of the bed. At the height of the flow distributor 38, the superficial liquid velocity must be higher than the minimum fluidization velocity for said particle size, and the bed expansion in the conical expansion section 39 (at the height of the flow distributor) computed for the average particle size should be in the range from 50 to 150 %, typically approx. 70 %. The cone angle (full angle) should be 1 ° to 45°, typically approx. 15°.

Figure 4 shows an alternative embodiment of the above-described implementation, this embodiment having a flow distributor 47 placed in the reactor 41 between the fluidized bed 45 and the fixed bed 46. As shown in the diagram of the figure, this embodiment has an drawoff nozzle 48 for liquid recycling adapted above the fluidized bed 45. The recycling line has a heat exchanger 44, which can be used for cooling the liquid mixture if necessary. The product mixture received from the fixed bed 46 is, however, routed entirely as such to further processing.

The most significant benefit of the above-described reactor with sidestream drawoff in MTBE/ETBE/TAME/TAEE processes is that the reacted product flow can be divided into two portions, namely the one recycled back to the fluidized-bed section 45 of the reactor in order to maintain the fluidized state and the one permitted to pass through the fixed-bed section 46. In terms of reaction

performance, this approach achieves a slightly higher conversion than such a reactor in which the entire volume of liquid passed through the reactor is recycled to maintain fluidization. The achievable gain in the conversion percentage is typically in the order of 5 to 20 % -units depending on the processed product, fraction of catalyst charge in the fixed bed and reaction temperature in the fixed bed. As the catalyst escapes the reactor along with the drawoff sidestream, it must be separated from the liquid flow with the help of, e.g., a cyclone (not shown) or a filter connected to the drawoff nozzle in order to reduce the slurry concentration of the sidestream to a sufficiently low level (5 to 40 %) for pumping with conventional pumps.

A reactor equipped with a sidestream drawoff is operated as follows:

- the catalyst is loaded in the reactor 41,

- the superficial liquid velocity in the reactor is slowly increased above the minimum fluidization velocity with simultaneous recycling of the liquid through the reactor up to a velocity achieving the expansion of a desired fraction of the catalyst charge into the fixed-bed section 46 in the reactor upper section,

- the superficial liquid velocity is reduced to a desired steady-state

fluidization level of superficial velocity, whereby a flow distributor 47 placed in the reactor upper section provides a catalyst-void region between the fixed-bed section 46 and the fluidized-bed section 45. The sidestream drawoff is located in this region.

- recycling is commenced via the sidestream drawoff nozzle.

To aid the reaction kinetics, a possibility of reducing the bed temperature with the progress of the reaction is desirable. According to the methods of conventional technology, cooling in a fluidized-bed reactor can be achieved by providing the bed with heat transfer tubes. However, as the fluidized bed is effectively mixed, this approach fails to achieve a temperature gradient in the desired direction.

A further desirable property is a possibility of minimizing the fraction of catalyst to be replaced during catalyst change. In the above-mentioned etherification processes, catalyst deactivation is mainly caused by catalyst poisons entering along the reactant feed lines. Therefore, the fraction of catalyst directly encountering the feed stream is advantageously limited. Isolating the catalyst charge in a fluidized-bed section in different fractions requires the use of flow distributors within the bed. However, if the perforated separator plate placed in the bed has a too low pressure drop, internal recirculation of the catalyst particles past the separator screen cannot be prevented. Both advantageous properties mentioned in the two paragraphs above can be obtained simultaneously by means of a heat exchanger which is located within the fluidized-bed section or after it, whereby a separate flow distributor is not required. Figs. 5 and 6 show two alternative heat exchanger constructions. The reactor arrangement shown in Fig. 5 corresponds to that illustrated in Fig. 3 with the exception that a tubular heat exchanger 57 is adapted within the fluidized-bed section 54 of the reactor 51. The process flow passes through the tubes of the heat exchanger 57 which also act as intermediate distributor in the fluidized-bed section. Cooling water is pumped through the cool side of the heat exchanger. Analogously, the reactor arrangement shown in Fig. 6 corresponds to that illustrated in Fig. 4 with the exception that a tubular heat exchanger 67 is adapted within the fluidized-bed section 65. The function of the heat exchanger is controlled by a temperature controller 68.

The final step of the reaction is preferably performed using a low degree of mixing. As is evident from the above-discussed, the degree of mixing can be lowered according to the invention by reducing the superficial liquid velocity below the minimum fluidization velocity, which is attained by either using a conically expanded reactor top section as shown in Figs. 3 to 6, or alternatively, arranging a sidestream drawoff for a portion of the reactant flow after the fluidized-bed section, whereby the reactor must be provided with a flow redistributor.

The latter preferred embodiment of the invention is shown in Fig. 7. This embodiment is principally analogous to that illustrated in Fig. 2 with the exception that a fixed-bed reactor 76 is adapted above the fluidized-bed reactor tube 75 operated with recycling, both beds being arranged into a single contiguous reaction space. Liquid flow into the fixed bed is limited by diverting at least a portion of the reaction mixture exiting the fluidized bed to a recycling loop, where the mixture is cooled when necessary prior to pumping it to the reactor inlet nozzle. Above the sidestream drawoff is arranged a tabular heat exchanger 77, and further above this, a flow distributor 78. Also the entry side of the fluidized bed is provided with a flow distributor 78.

In addition to the above-described characteristics, it must be noted that the arrangement provides easy recharging and discharging of particulate matter from the fluidized bed as the fluidized particles behave like a fluid. Consequently, catalyst recharging and discharging can be implemented using an arrangement illustrated in Fig. 8 for example.

Fig. 8 shows a system which facilitates run-time catalyst change during manufacture of tertiary ethers. With reference to Fig. 3, the apparatus comprises a reactor 81 incorporating a feed pump 82 and heat exchangers 83 and 84, the latter of which adapted to the recycle loop.

In a fluidized state the catalyst acts as a flowing liquid, whereby it can be discharged from the reactor 81 simply by opening the nozzle 85 which is located below the upper surface of the fluidized bed. A fluidized-bed reactor is rarely discharged fully, but rather, a seeding bed is left in the reactor. Further, the discharge point is located as close to the reactor flow distributor 86 as possible, since such an arrangement assures the discharge of large particles probably concentrated there. To achieve a controlled discharge, the amount of discharged catalyst must be controllable. In the simplest manner this takes place with the help of a discharge container 87 that acts as a restriction for the catalyst charge limiter pipe which is routed to the interior of the discharge container, whereby the maximum charge is limited to the lower end of said limiter pipe. The size of the discharge container is determined by the volume of desired discharge fraction of the catalyst charge; however, the container volume may not exceed the total volume of the catalyst contained in the reactor in fixed-bed form. The discharge container 87 is placed either level with the reactor catalyst discharge nozzle 85 or below it.

The discharge procedure is as follows:

- The container 87 is empty and purged with nitrogen. The pipeline 88 to the top of the reactor 81 is filled with the liquid. All valves of the container are closed.

- To avoid purging of the reactor with inert gases, the container 87 is filled with the liquid recycled through the reactor by opening the valve 89. The container pressure is kept above the vapour pressure of the liquid with the help of a pressostat 90 that controls a valve 91. The reactor is filled with caution.

- When the level sensor switch 92 indicates the container 87 to be filled with the liquid, the valves 93 and 94 are closed.

- The valves 95 and 96 are opened, whereby the catalyst can flow into the container. If the valves are kept open sufficiently long (1 to 2 min), the container is filled up to the level of the limiter pipe with the catalyst and the liquid. The correct transfer sequence is assured by level sensors.

- The valves 95 and 96 are closed.

- The valve 91 is opened to allow the container pressure fall below the vapour pressure of the liquid, whereby the liquid is evaporated away. The container must have a heating jacket to prevent excessive cooling. If the volatile hydrocarbons are not collected using a recovery system, but rather, are routed to the flare for combustion, the valves 91 and 94 can be combined.

- The valve 97 is opened, whereby nitrogen is purged through the catalyst bed, thus evaporating any hydrocarbon residues (during a few hours typically).

- The valve 98 is opened, whereby the spent catalyst falls into, e.g., a bin for transport away.

- The valves 91, 97, 98 are closed.

- The procedure can be repeated from the beginning. Recharging of the reactor with catalyst is easy. Catalyst recharging in batches can be performed with the help of a feed container 99 located above the upper surface of the reactor catalyst bed (the minimum requirement being that the container bottom is located above the reactor flow distributor 86). The volume of the container 99 is designed according to the desired maximum recharge batch. The maximum size of the container may not exceed the catalyst volume in the reactor in fixed-bed form. The recharge procedure is as follows:

- The container 99 is empty and purged with nitrogen. All valves of the container are closed.

- The container is filled with the catalyst via the manhole. During the fill-up operation the container 99 is purged with nitrogen via an inlet valve 100, while an outlet valve 101 is simultaneously held open to allow the nitrogen to escape from the container.

- The manhole is closed.

- If the catalyst is slurried in water, the container bottom valve 102 is

opened and the water is drained, after which the valve is closed.

- The valves 100 and 101 are closed.

- To avoid purging of the reactor with inert gases, the container 99 is filled with the liquid recycled through the reactor 81 by opening the valve 103. The container pressure is kept above the vapour pressure of the liquid with the help of a pressure control valve 104, which controls the valve 105. The reactor is filled with caution.

- When the level sensor switch 106 indicates the container 99 to be filled with the liquid, the valves 103 and 105 are closed.

- The valves 107 and 108 are opened, whereby the catalyst can flow into the reactor. The container and the recharge pipe are flushed with the liquid via the valve 106. The correct transfer sequence is assured by level sensors.

- The valves 107 and 108 are closed. The container is now filled with the liquid. The container can be dumped into the reactor with the help of nitrogen pressure if nitrogen is available at sufficient pressure. If not, the container contents can be purged to the flare for combustion.

- The procedure can be repeated from the beginning. The invention is elucidated in greater detail with the help of the following application examples.

Example 1 (reference)

Manufacture of TAME in a fluidized-bed reactor

With reference to Fig. 2, the apparatus employed was a reactor tube in which the methanol was reacted with the isoolefin mixture in a conventional fluidized bed. The reactor was provided with recycling of the products.

Reactor inner diameter 154.1 mm

Reactor height 7500 mm

Catalyst type Amberlyst 15 cation exchange resin

Catalyst quantity 66 1 (not in fluidized state)

Fluidization characteristics with the reaction mixture:

Bed expansion Superficial liquid velocity v

(BE) % m/h

0 27

10 47

30 87

The expansion of the bed can be estimated with the help of an equation derived from Eq. 11:

BE/100 % = 0.0050 · v/(m/h) - 0.135

Feed composition

Methanol 7.96 wt.-%

C_4- 9.16 wt.-%

2-methyl-2-butene 10.80 wt.-%

2-methyl-1-butene 5.73 wt.-% C₅ compounds, other 52.74 wt.-%

C₆₊ 13.61 wt.-%

Reactor fresh reactant feed rate 117 kg/h

Reactor total feed rate 695 kg/h recycling inclusive

Reactor operation temperature 65 °C

Achieved conversion of isoamylenes (2-methyl-2-butene + 2-methyl-1-butene = total isoamylenes) into TAME 51.4 %

Example 2

Manufacture of TAME in a combination fluidized/fixed-bed reactor tube With reference to Fig. 7, TAME was produced using equipment in which the apparatus employed in Example 1 was complemented with a heat exchanger (once-through-type on the hydrocarbon side), and above that, a 4500 mm high pipe of 154.1 mm ID with a perforated plate-type flow distributor 'placed at the lower section of the pipe. The volume of catalyst fraction in this pipe section during normal operation is maintained at 70 1. The reaction mixture was cooled to 50 °C by the heat exchanger. The recycling sidestream is drawn off from below the heat exchanger. The other reaction conditions were kept the same as in Example 1. The obtained conversion of isoamylene (into TAME) was 60.7 %. Example 3

Production of TAME in a reactor with two expanded sections

With reference to Fig. 3, the reactor employed in the example had a fixed bed arranged above the fluidized bed with the help of a 30 % expansion section adapted in the reactor, whereby said expansion section accomplished the reduction of the liquid mixture superficial velocity below the minimum fluidization velocity. A portion of the product flow was recycled back to the inlet nozzle of initial reactants.

Reactor diameter at flow distributor 650 mm

Cone angle of reactor bottom 30°

Reactor diameter along fluidized-bed section

(section of 30 % expansion) 900 mm

Height of fluidized-bed section (section of 30 % expansion) 25500 mm Catalyst quantity in fluidized-bed section (not in fluidized state) 13.2 m³

Cone angle of conical section after fluidized-bed section 30°

Reactor diameter at conical section top 2000 mm

Catalyst quantity in fixed-bed section (not in fluidized state) 15 m³

Height of fixed-bed section (section of 0 % expansion) 5000 mm

Total height of catalyst bed in reactor

(can be computed using Eqs. 4, 13 to 15, 18 and 19) 31.0 m

Height of catalyst bed in stopped reactor 29.8 m Catalyst

Same as in Example 1.

Feedstock composition:

Same as in Example 1.

Reactor fresh reactant feed rate 16,25 t/h Reactor total feed rate 29 t/h

Reactor feedstock temperature 65°C Obtained total conversion of isoamylenes into TAME 52.1 %

Example 4

Manufacture of TAME in a combination fluidized/fixed-bed reactor equipped with a heat exchanger

The apparatus and the reactant mixture were the same as in Example 3 with the exception that the fluidized-bed section was provided at 15000 mm height level from the flow distributor with a 3000-mm long heat exchanger comprising 1160 pcs. of OD ¾ " heat exchanger tubes (cf. Fig. 5). With the help of the heat exchanger, the reaction mixture otherwise identical to that used in Example 3 was cooled to 50 °C.

Obtained total conversion of isoamylenes into TAME 53.9 % Example 5

Manufacture of TAME in a combined fluidized/fixed-bed reactor equipped with a heat exchanger and recycling of intermediate products

The apparatus and reaction mixture were the same as in Example 4 with the exception that the bottom part of the upper conical section was provided with a flow redistributor and the sidestream drawoff was adapted to the reactor below said distributor (cf. Fig. 6).

Obtained total conversion of isoamylenes into TAME 57.3 % Example 6

Manufacture of ETBE

The example was carried out using the apparatus described in Example 5, while the reaction mixture was as follows:

Hydrocarbon feedstock: Ethanol feedstock

Isobutene 22.33 wt.-% Ethanol 99.38 wt.-%

Isobutane 31.58 ETBE 0.16

n-butane 10.03 H₂O 0.46

1-butene 12.82

Trans-2-butene 12.94

Cis-2-butene 9.27

C₅₊ 1.03

Feed rate 10 t/h Feed rate 1.5 kg/h

The total feed rate to the reactor was 30 t/h and the feedstock temperature was 50 °C. In these conditions the conversion of isobutene into ETBE was 76.9 % .

Example 7

Manufacture of MTBE Apparatus, hydrocarbon feed and reaction conditions as in Example 6, while the alcohol was changed to methanol and processed at a feed rate of 1.1 t/h.

Obtained conversion of isobutene to MTBE 81,4 %

Claims

Claims:

1. A process for producing tertiary alkyl ethers from isoolefins or mixtures thereof and from at least one lower aliphatic alcohol, according to which process - the reactants are formed into a liquid mixture which is fed through at least one first reaction zone (35; 45; 55; 65; 75) incorporated into a reactor vessel, said zone containing catalyst particles made from an acid cation exchange resin which form a first reaction bed, while the liquid mixture superficial velocity is kept sufficiently high so that said first reaction bed comprised of the catalyst panicles expands at least to some extent during the throughflow of the liquid mixture,

- the liquid mixture passing the first reaction zone is routed through at least one second reaction zone (36; 46; 56; 66; 76) formed by a second, at least substantially fixed bed, said second bed containing similar catalyst particles as said first bed, and

- if necessary, the product obtained from said second reaction zone is routed to further processing and separation of the product components, c h a r a c t e r i z e d in that

- said second reaction zone (36; 46; 56; 66; 76) is arranged in the same reactor vessel as said first reaction zone and

- the mixture passing said first reaction zone (35; 45; 55; 65; 75) is routed as an upflow through said second reaction zone using a superficial liquid velocity smaller than the minimum fluidization velocity thus forming an at least essentially fixed reaction bed in said second reaction zone.

2. The process in accordance with claim 1, wherein the liquid mixture superficial velocity is reduced after passing the first reaction zone (35; 45; 55; 65; 75).

3. The process in accordance with claim 1, wherein the liquid mixture superficial velocity is reduced to a value which is approx. 20 to 40 %, preferably approx. 30 to 35 %, of the superficial velocity prevailing in the first reaction zone (35; 45; 55; 65; 75).

4. The process in accordance with claim 2 or 3, wherein to the end of reducing the liquid mixture superficial velocity, the cross-sectional area of the second reaction zone (36; 46; 56; 66; 76) is larger than that of the first reaction zone.

5. The process in accordance with any one of the claims 1 to 4, wherein at least a portion of the mixture passing the first reaction zone (35; 45; 55; 65; 75) is drawn off before the second reaction zone and recycled to combine with fresh reactant feed to the first reaction zone.

6. The process in accordance with any one of claims 1 to 5, wherein a zone (39) with a high degree of mixing is arranged prior to the entry of the reaction mixture into the first reaction zone.

7. The process in accordance with claim 6, wherein the expansion factor of the first reaction zone is at least approx. 50 % on the entry side (39) of the reaction zone and preferably at least approx. 10 % on its exit side, while the expansion factor of the second reaction zone is smaller than 5 %.

8. The process in accordance with any one of claims 1 to 7, wherein the reactant mixture is cooled prior to its entry into the second reaction zone.

9. An apparatus for producing tertiary alkyl ethers from isoolefins or mixtures thereof and at least one lower aliphatic alcohol using a process in which the isoolefins and alcohols are mixed to form a reactant mixture and reacted with each other in the presence of a particulate cation exchange resin, said apparatus comprising

- a contiguous, elongated reactor vessel (31; 41; 51; 61; 71), which can be filled with the cation exchange resin used in the process, said vessel having an essentially vertically aligned longitudinal axis and incorporating an inlet nozzle (30A; 40A; 50A; 60A; 70A) for reactant mixture feed and an outlet nozzle (30B; 40B; 50B; 60B; 70B) for removal of products, c h a r a c t e r i z e d in that - the reactor vessel is provided with a lower section, in which a fluidized- bed section (35; 45; 55; 65; 75) can be formed, and an upper section, in which a fixed-bed section (36; 46; 56; 66; 76) can be formed,

- the reactor vessel lower section (35) is joined to reactor vessel upper

section (36) via an expansion section (37), which is shaped as a truncated cone with a cone angle of 1° to 30°,

- the volume of the upper section amounts to 30 to 80 % of the total volume of the reactor vessel, and

- the inlet nozzle (30A; 40A; 50A; 60A; 70A) for reactant mixture feed is fitted to the reactor vessel lower section and the outlet nozzle (30B; 40B; 50B; 60B; 70B) for removal of reaction products is fitted to the reactor vessel upper section, respectively.

10. The apparatus in accordance with claim 9, wherein the fixed reaction bed (36; 46; 56; 66; 76) can be formed by reducing the superficial velocity of the liquid flowing through the reactor in the reactor vessel upper section to a value below the minimum fluidization velocity and the catalyst contained in the fixed bed can be changed by temporarily reducing the superficial liquid velocity so much as to allow the catalyst particles contained in the fixed bed to fall by gravity to the lower section of the reactor, wherefrom the spent catalyst can be unloaded as necessary and replaced by new catalyst.

11. The apparatus in accordance with claim 9, wherein the volume of the upper section is about 40 to 80 %, preferably about half of the total volume of the reactor vessel.

12. The apparatus in accordance with claim 9, wherein the cross-sectional area of the reactor vessel increases above the expansion section so much as to permit the superficial velocity of the liquid upflow from the lower section to the upper section to fall below the minimum fluidization velocity of the particle mixture contained in the upper section.

13. The apparatus in accordance with claim 12, wherein the volume of the lower section is less than the total volume of the particulate cation exchange material.

14. The apparatus in accordance with claim 9, wherein the expansion section (37) is formed by a collar section (37) with a cone angle of 5° to 15°.

15. The apparatus in accordance with claim 12, wherein the cross-sectional area of the reactor vessel at the upper section is approx. 2- to 10-fold the reactor vessel cross-sectional area at the lower section.

16. The apparatus in accordance with any one of claims 9 to 15, wherein an inlet nozzle (48) of a recycling line is fitted essentially above the lower section of the reactor vessel, said inlet nozzle permitting recycling of at least a fraction of the product mixture passing through the reactor vessel lower section back to the inlet nozzle of the initial reactants.

17. The apparatus in accordance with any one of claims 9 to 16, wherein the lower and upper sections of the reactor vessel are separated by a liquid flow distributor (47).

18. The apparatus in accordance with claim 9, wherein the inlet region at the lower section of the reactor vessel is formed into a conically shaped expansion section (39) in order to achieve a zone of a high degree of mixing.

19. The apparatus in accordance with claim 9, wherein the lower section of the reactor vessel is provided with a heat exchanger (57; 67; 77) adaptable to cool the reaction mixture passing the fluidized bed prior to the entry of the mixture into the fixed bed.

20. The apparatus in accordance with any one of claims 9 to 19, wherein the inlet nozzle of the reactor vessel is connected to a flow control element of the initial reactants, said flow control element being suitable for the control of the liquid feed flow rate into the reactor vessel.

21. The apparatus in accordance with claim 20, wherein the lower section of the reactor vessel is provided with loading (108) and unloading (95) nozzles of cation exchange resin.

22. A prereactor used in the production of MTBE, ETBE, TAEE or TAME, characterized in that it comprises an apparatus in accordance with any one of claims 9 to 21.

23. An etherification reactor used in the production of TAME, characterized in that it comprises the apparatus in accordance with any one of claims 9 to 21.

24. A method for operating the apparatus in accordance with any one of claim 9 to 23, characterized by

- filling the reactor vessel (31; 41; 51; 61; 71) at its lower section with a charge of cation exchange resin particles sufficient for forming two reaction beds.

- feeding the liquid mixture of initial reactants to be reacted with each other is fed via the inlet nozzles (30A; 40A; 50A; 60A; 70A) of the reactor vessel to the lower section of the reactor vessel,

- adjusting the liquid mixture superficial velocity so high that the cation exchange resin particles form a fluidized bed (35; 45; 55; 65; 75) in the vessel lower section and the excess fraction of the catalyst charge is transferred along with the liquid mixture flow to the upper section of the reaction vessel,

- keeping the liquid mixture superficial velocity in the upper section of the reactor vessel below the minimum fluidization velocity in the order to form a fixed bed (36; 46; 56; 66; 76) in the upper section, - maintaining the fluidized bed in the lower section of the reactor vessel and the fixed bed in the upper section, respectively, during the reaction run of the initial reactants. and

- if necessary, removing a fraction of the cation exchange resin contained in the fluidized bed during the reaction run and replacing it by new or regenerated resin.

25. The method in accordance with claim 24, wherein the catalyst is changed by

- opening a catalyst unload nozzle (95) at the lower section of the reactor

(81),

- keeping the superficial liquid velocity above the minimum fluidization velocity until a desired fraction of the catalyst is removed from the reactor,

- reducing the liquid flow rate in the reactor at least to a level that allows the catalyst contained in the upper section of the reactor to fall down,

- increasing again the superficial liquid velocity in the reactor above the minimum fluidization velocity so expanding a desired fraction of the catalyst into the fixed-bed section at the upper section of the reactor vessel, and

- feeding a fresh catalyst replacement into the reactor and reducing the

superficial liquid velocity to the desired fluidization velocity level.