DD230545A3 - PROCESS FOR THE PRODUCTION OF LOW OLEFINES AND FLAVORS - Google Patents

Abstract

The invention describes a process for the preparation of lower olefins, preferably ethylene, and monoaromatics by catalytic conversion of methanol and hydrocarbons. The aim is to demonstrate a technically feasible process that couples the exothermic methanol conversion with the endothermic cleavage reaction of the hydrocarbon such that both feedstocks contribute to increasing olefin selectivity and better control of the exothermicity is possible. According to the invention, mixtures of methanol with liquid and / or gaseous hydrocarbons, u. a. those of the gasoline to Gasoelsiedebereiches in Mengenverhaeltnis 10: 1 to 1: 5 of high-silica zeolite-rich zeolite catalysts with channel structure (modulus 35, effective pore diameter 0.4 to 0.6 nm) at temperatures of 600 to 800C, at a load of 0 , 5 to 6.0 v / vh, at a ratio of fed-in steam to the feedstock of 0.5 to 10: 1 and reacted at a pressure of 0.1 to 2.0 MPa.

Description

Field of application of the invention

The invention relates to a process for the preparation of lower olefins, preferably ethylene, and aromatics by catalytic conversion of methanol and hydrocarbons.

Characteristic of the known technical solutions

The production of olefinic hydrocarbons by dehydration of methanol is known in principle. As early as 1962, Mattox demonstrated the formation of olefins from methanol on NaX zeolites (US Pat. No. 3,036,134). Significant progress in methanol conversion to lower olefins has been achieved by P.B. Weisz (Chimia33 [1979] 154) with the use of a new zeolite catalyst characterized by a high modulus.

On this basis, DLKing et al. (Catal. Rev. Sei. Eng. 23 [1981] 248) and YCHu (Hydrocarbon Processing 1983, 93) developed the conditions to be achieved for the achievement of high yields, especially of low molecular weight olefins in methanol conversion. Thereafter, both the use of a high space velocity and the dilution of the feed methanol as well as the attenuation of the surface acidity of the zeolite by addition of water and the increase of the shape selectivity of the catalysts by modification are of fundamental importance. Furthermore, mild reaction conditions favor the formation of the olefins, so that for the methanol conversion to zeolitic catalysts generally temperatures below 500 ° C are given. As a preferred temperature range 260 to 450 ⁰ C called (US-PS 4049573, 4088706). In this aspect, the use of the following types of catalyst is proposed:

- aluminosilicate zeolites (US-PS 4079095, 4079096)

Manganese-doped aluminum silicates (EP-PS 0021 330, 0023303, US-PS 4247731)

- Catalysts with a pore size <0.6nm, z. B. Chabazite exchanged with rare earths (Hydrocarbon Processing 1983, 93)

- Modified HZSM-5 systems with phosphorus compounds or antimony trioxide (US Patent 3972832, GB-PS 1 528 674) or by ion exchange with cations of a radius above 0.1 nm (US Patent 4093543).

The variety of known technical teachings for modifying the zeolite catalysts (DE-OS 2542230, 2633881, US-PS 3911 041, 3979472,4049573,4049735,4066714, 4231 899, EP-PS 0073482, 0073483) as well as for the development of zeolites with a defined , in particular characterized by the crystallinity geometry (DE-OS 3139355, US-PS 4025571, 4025572,4148835,4172856,4278565,4289607, 4361715) leaves the skilled in the art essentially processed, the most promising main directions in the methanol conversion to lower olefins recognize. In addition, however, the use of so-called inert diluents for use of methanol has attracted attention. In this sense, especially hydrogen, helium, nitrogen, carbon dioxide, lower alkanes C ₁ to C ₃ or exhaust gases (DE-OS 2935863), water vapor and paraffins C ₁ to C ₆ (US-PS 4025575,4025576,4083888,4083889) , Anhydrous thinner (US Patent 4093888) described and the addition of at least 50 wt .-% of water to the starting product (DE-PS 2827385) proposed.

In addition to these two developments directed towards the catalyst system or to the medium to be converted, however, several procedural arrangements are also proposed with the aid of which a controllable and thus controllable temperature control can be achieved during the exothermic methanol dehydration (US Pat. No. 3,894,104, 4,051,449,413,884,423,831,3). 4328384).

From the wealth of known technical teachings, however, no such solution can be found that satisfies in its entirety in the implementation of methanol conversion to lower olefins both the procedural problems that are caused mainly by the high exothermicity, as well as a satisfactory in itself Olefin yield gives the expected C ₂ , C ₃ , C ₄ olefins in a selectivity ratio shifted in favor of ethylene. For the effective conversion of methanol to olefins and the associated control of the exothermic is usually a high technological effort to realize that consists for example in a two-stage reaction (US-PS 3894104), in fixed bed arrangements of the catalyst, the introduction of intercooling (US-PS 4138440) and in the case of the reaction in the fluidized bed for heat dissipation makes the installation of complex and complicated internals required and beyond high demands on the abrasion resistance of the catalyst is (US-PS 4138440). The proposed according to US-PS 4238631 to maintain the catalyst temperature high catalyst feed ratio of 10 and 20: 1 inevitably involves only a small and therefore uneconomical space-time yield. The further associated with the two-stage reaction indirect cooling of the intermediate dimethyl ether-methanol-water (US-PS 4052479), doing additionally occurring separation problems (DE-PS 3135618) and described in DE-PS 3118954 direct dependence of the olefin selectivity of methanol conversion - the higher the methanol conversion, the smaller the olefin yield - point to further shortcomings of this procedure.

The use of inert diluents to achieve increased olefin selectivity requires at a methanol conversion of between 40 and 90% by weight and at an ethylene yield of 18% by weight dilution ratios for anhydrous diluents between 2 and 20 mol / mol of methanol and between 3 and 10 mol of steam together an inert diluent whose dew point is not higher than 40 ° C, per mole of methanol (US Patent 40838884083889). Since the reaction conditions used in the use of hydrocarbons as diluents temperature maximum _ to 550 ⁰ C - cause no satisfactory pyrolysis to C ₂ olefin (US-PS 4172856) results in this way also a very small space-time yield. However, due to the nature of the decomposition of CO / H ₂ in methanol (see RH Ewell, Ind. Eng. Chem732 [1940] 149), temperature increase in the temperature range of the preferred hydrocarbon pyrolysis was not taken into account because the adjustment promotes this balance by different catalytic influences (see: Berdsch., former Ges., 35, 1055; Compend., 148, 1735; Liebigs, Ann. 318, 146, Angew. Chem., 1936, 485).

Object of the invention

It is the object of the invention to provide a technically feasible process for obtaining olefinic hydrocarbons of low molecular weight in addition to monoaromatics from mixtures of oxygen-containing organic compounds and hydrocarbons, with the exothermic reaction of the oxygen-containing organic compounds with the endothermic cleavage reaction of the hydrocarbons for better control of Exothermic be coupled and contribute both reactants to increase the olefin selectivity, in particular in favor of ethylene formation.

Explanation of the essence of the invention

The object is achieved by means of a process for the preparation of lower olefins, preferably ethylene, and aromatics, according to the invention mixtures of methanol with liquid and / or gaseous hydrocarbons in a ratio of 10: 1 to 1: 5 in the presence of a high-siliceous zeolite with channel structure at temperatures Range of £ 00 to 800 ° C, preferably in the range of 650 to 74O ⁰ C, at a load of 0.5 to 6.0v / vh, with a ratio of fed steam to the feed of 0.5 to 10: 1 and be reacted at a pressure of 0.1 to 2.0 MPa. Preference is given to using mixtures of methanol with hydrocarbons of the gasoline to gas oil boiling range; in this case, instead of the methanol, it is also possible to use the dimethyl ether or a mixture which optionally contains water and consists of methanol and dimethyl ether.

As a high-silica zeolite having a channel structure, a system with an effective pore diameter of 0.4 to 0.6 nm and a modulus of> 35 has been proven.

Further features of the solution according to the invention consist in that the heat of reaction liberated during the methanol conversion is consumed for the endothermic process of hydrocarbon splitting and water vapor forming in situ is used for the partial or complete substitution of the process steam for the hydrocarbon splitting in the methanol conversion.

The process of the invention should be carried out in particular in such a way that both superheated methanol vapor and superheated hydrocarbon are passed simultaneously over the high-silica zeolite Kataiysator, the material flows are advantageously to be dimensioned so that the exothermic and endothermic enthalpies of the ongoing reactions largely compensate and replaced the reaction water formed as much of the required process water. To initiate the desired reaction processes, the zeolite catalyst requires high thermal stability and good water vapor resistance. Notwithstanding the behavior commercial zeolites which Matalysatoren used in the invention convert completely surprising even at temperatures above 600 ⁰ C, preferably about 700 ⁰ C and in the presence of fissile hydrocarbons of methanol or a made of Kethanol in a precursor dimethyl ether or consisting of methanol and dimethyl ether Mixture with high selectivity in lower olefins in addition to aromatics without appreciable formation of carbon monoxide and with only relatively low methane formation. In addition, the catalyst causes a likewise unexpected increase in the degree of conversion in the decomposition reaction of the hydrocarbons to the expected lower olefins.

The technical means required for the realization of the solution according to the invention comprise in each case an evaporator for the separately fed, to be converted or to be split substances aqueous methanol and hydrocarbon mixture, which are then combined for common overheating in a corresponding aggregate and then at space velocities above 0.5v / vh , preferably around 5v / vh, based on both the methanol and the hydrocarbon, and at the preferred temperature range of 650-740 ° C, at normal or only moderately elevated pressure over the catalyst. The catalyst zone preferably consists of a fixed bed, but it can also be designed as a fluidized or fluidized bed. The reaction mixture leaving 650 to 700 ° C hot reaction mixture is preferably cooled indirectly indirectly in a quench cooler using the sensible heat to about 350 ° C and after passing a direct oil and water wash to room temperature. The reaction mixture separates into under normal conditions liquid and gaseous hydrocarbons. The gaseous reaction mixture is decomposed in a subsequent gas separation plant in the usual way in the target products. The resulting during the process water is partially removed from the steam stripping and can be used, for example, as industrial dew water.

By changing the methanol-hydrocarbon ratio within the defined limits as well as by a water addition to the feed methanol and also by varying the superheater outlet temperature, the process control can be thermally controlled and the product spectrum can be influenced.

The highly silica-rich zeolite catalyst having channel structure used in the invention, an effective pore diameter of 0.4 to 0.6 nm and a modulus of> 35 has e.g. a mean particle size of 15nm and lattice constants for a = 2.016; b = 1.994; c = 1,346nm. By cation exchange of the catalyst due to doping with metals, for example with cobalt, thorium, ruthenium, vanadium, the selectivity ratio of the resulting target products can also be varied within certain limits.

A reduction of the methanol hydrocarbon partial pressure in the catalyst zone by addition of steam counteracts the unwanted carbon deposition on the catalyst and prolongs its transit time. The regeneration of the catalyst is carried out by burning off the deposited carbon with a vapor / air mixture at temperatures in the range of 700 and 800 ⁰ C.

The economic and technical advantages resulting from the use of the method according to the invention can be summarized as follows:

- Joint methanol conversion and hydrocarbon cleavage in a process stage for the formation of lower olefins with increased ethylene content.

- Thermal advantages by using the energy bound in methanol to C-C cleavage of the hydrocarbon.

- Course of the catalytically initiated hydrocarbon pyrolysis at 100 to 150 ^c C lower operating temperatures, compared with the thermally initiated pyrolysis.

- The catalyst stage can be integrated instead of the tubular reactor in a conventional gasoline splitter.

- For fission separation no fundamentally new technology is required.

The process is not tied to the provision of water and energy for process steam generation, as in exclusive hydrocarbon pyrolysis.

- The by-produced reaction water can be used for example for steam generation.

embodiments

Examples 1 to 26 shown in Tables 1 to 8 represent test results which were obtained on an electrically heated laboratory reactor made of quartz (0 40 mm).

The catalyst used should be characterized as follows:

High-silica-rich zeolite with channel structure

Effective pore diameter 0.5 to 0.56 nm

Molar ratio SiO ₂ IAI ₂ O ₃ = 35

Average particle size 15nm

Lattice constant a = 2.016; b = 1.994; c = 1,346nm

The experiments were carried out in such a way that the liquid feed products in a predetermined ratio initially added to an inert evaporator layer, superheated after evaporation and then the catalyst zone (about 20 ml) were fed together. The resulting reaction products were then quenched indirectly, the separated liquid product separated into an aqueous and a hydrocarbon phase and the reaction gas also obtained volumetrically determined. All reaction products were analyzed by gas chromatography and balanced. The examples prove, in addition to the solution according to the invention for comparison purposes, both the catalytic conversion of methanol / water and hydrocarbon / water mixtures in the temperature range between 680 and 740 ^c C.

The individual test parameters, such as reaction temperature, feedstock / water mass ratio, test duration and catalyst loading, and the yields of the products obtained in% by mass are shown in Tables 1 to 8.

The yields given in Tables 1 to 8 for CO and CO ₂ are based on the input product sum of methanol and hydrocarbon equal to 100.

Table 1 (Examples 1 to 3) Methanol Methanol - Methanoi Feed products Water Water Water

Temperature ( ⁰ C) 680 710 740

Methanol / water mass ratio 1: 1 1: 1 1: 1 catalyst loading (v / vh) 5.5 5.7 5.4

With methanol 3.1 3.2 3.0

- with water 2.4 2.5 2.4

Yields (% by mass) * * *

Hydrogen "1.7 4.9 6.8

Methane. 23.5 26.5 20.9

Ethylene 11.8 10.2 5.0

Propylene 7.0 '6.3 1.8

Butadiene 0.7 1.0 0.3

Σ cracking gas <C ₄ 69.1 95.2 98.8 Benzene toluene xylene 0.4 2.9 15.9 0.2 0.1 0.1 0.1 Pyrolysis gasoline <200 ° C 23.1 4.8 1.2 Pyrolysis oil> 200 ° C 7.8 - - CO CO ₂ 16.6 2.1 38.9. 2.8 58.6 3.7 Methanol (% by mass) 84.0 69.5 72.0

based on CH ₂ content of the reacted methanol

Table 2 (Examples 4 to 7) use products

n-decane water

n-decane water

n-decane

methanol

water

n-decane

methanol

water

Temperature ( ⁰ C) feedstock / water mass ratio catalyst load (v / vh)

- with n-decane - with methanol

- with water

680

1: 0.7 without catalysis

680 1: 1.4

6 3

680 1: 0.7: 0.5

680 1: 0.5: 0.5

2.9 1.7 0.6 0.6

Yields (% by mass)

hydrogen

methane

ethylene

propylene

butadiene

Table 3 (Examples 8 to 11) use products

D 0.3 2.7 10.5 5.2 0.4

n-decane water

0.4

3.6 17.8 18.8

0.6

n-decane water

0.6

5.5 17.8 17.2

0.6

n-decane

methanol

water

2)

0.6

6.6

20.2

19.9

0.9

Σ cracking gas <C ₄ 23.3 53.2 55.0 63.7 Benzene toluene xylene 0.9 0.5 0.6 0.3 0.2 0.1 0.5 1.4 1.7 0.6 1.4 1.5 Pyrolysis gasoline s200 ⁰ C 72.3 43.8 42.2 34.1 Pyrolysis oil> 200 ° C 4.4 3.0 2.8 2.2 CO CO ₂ 0 0 0 0 1.9 0.2. 2,3 0,2 Conversion n-decane (% by mass) conversion methanol (% by mass) 40 " 63 70 91 76 92

n-decane

methanol

water

Temperature ( ⁰ C) feedstock / water mass ratio catalyst load (v / vh)

- mitn-decane

- with methanol

- with water

710

1: 0.7 without catalysis

710 1: 1.4

5.1 2.5

2.6

710 1: 0.6: 0.6

5.1 2.6 1.5 1.0

710 1: 0.5: 0.5

2.5 1.7 0.6 0.6

Yields (% by mass)

hydrogen

methane

ethylene

propylene

butadiene

D 0.4 4.4 19.2 9.7 1.4

0.5

5.8 26.8 20.9

1.3

0.9

8.4 20.3 17.6

1.4

2)

0.8

9.7 24.7 18.6

1.7

Σ cracking gas <C ₄ 42.4 67.0 65.0 73.2 Benzene toluene xylene 1.1 0.5 0.1 0.7 0.7 0.3 0.9 1.9 2.1 1.0 1.4 1.4 Pyrolysis gasoline <200 ^° C. 53.5 31.0 34.2 24.9 Pyrolysis oil> 200 ° C 4.1 2.0 0.8 1.9 CO CO ₂ 0 0 0 0 3.8 0.2 , 4.6 0.2 Sales n-decane (% by mass) Sales methanol 60 75 79 95 85 93

refer to a used n-decane

- ο-

Table 4 (Example 12 to 15) use products

petrol

methanol

water

petrol

methanol

water petrol

Methanol water

petrol

methanol

water

Duration of test (h) Temperature (° C) Input product / water mass ratio Catalyst load (v / vh)

- with gasoline

- with methanol

- with water

680

1: 0.7: 0.7 2.6 1.2 0.8 0.6

680

1: 0.8: 0.8 2.4 1.0 0.8 0.6

1: 0.8: 0.8 2.5 1.1 0.8 0.6

680

1: 0.8: 0.8 2.4 1.0 0.8 0.6

Yields (% by mass)

hydrogen

methane

ethylene

propylene

butadiene

0.9

7.5

10.2

11.5

0.9

0.9

7.8

11.6

13.1

1.2 1.0

8.2

11.0

12.5

1.4

* Based on the gasoline used and the CH ₂ content of the methanol used Table 5 (Examples 16 to 18)

use products

Gasoline water

petrol

methanol

water petrol

methanol water

0.8

8.1

.10,4

15.5

1.4

Σ cracking gas <"C ₄ 43.6 49.0 48.5 50.5 Benzene toluene xylene 2.0 4.5 5.7 1.9 4.0 7.3 1.6 3.7 6.5 1.7 3.0 5.9 Pyrolysis gasoline <200 ° C 48.6 40.9 41.5 42.3 Pyrolysis oil> 200 ° C 7.8 10.1 10.0 7.2 CO CO ₂ 3.9 0.2 4.1 0.2 4,5 0,2 4.4 0.2 Methanol (% by mass) 85.7 86.4 86.4 · 87.1

Temperature ( ⁰ C) feedstock / water mass ratio catalyst load (v / vh)

- with gasoline

- with methanol

- with water

680 1: 1.5

5.3 2.6

2.7

680 1: 0.7: 0.7

5.3 1.6 2.5 1.2 680 1: 0.5: 0.5

3.0 1.6 0.8 0.6

Yields (% by mass)

hydrogen

methane

ethylene

propylene

butadiene

1)

0.2

2.1

5.8

8.2

0.6

2) 0.6 5.1 9.2 11.1 1.0 2)

0.6

5,6 11,0 12,5

0.9

Σ cracking gas <C ₄ 23.7 38.5 42.3 Benzene toluene xylene 1.3 2.6 2.2 1.9 2.6 4.3 2.2 4.2 4.7 Pyrolysis gasoline <200 ^° C. 69.0 54.0 49.6 Pyrolysis oil> 200 ° C 7.3 7.5 8.0 CO CO ₂ 0.1 0 2.4 0.2 2,2 0,2 Methanol (% by mass) - 77.5 85.8

¹¹ based on used gasoline

²¹ based on the gasoline used and the CH ₂ content of the methanol reacted

- ο - uwu ι τ

Table 6 (Examples 19 to 21)

use products

petrol petrol petrol methanol methanol water water water

Temperature ( ⁰ C) Feedstock / water mass ratio Catalyst loading (v / vh) with gasoline with methanol with water

710 1: 1.5

5.2 2.5

2.7 710 1: 0.7: 0.7

5.3 2.5 1.6 1.2

710 1: 0.4: 0.4

3.0 1.6 0.8 0.6

Yields (% by mass)

hydrogen

methane

ethylene

propylene

butadiene

0.5

5.5.14.14.9

1.8 2)

1.1

8.5

13.1

13.5

1.9

¹¹ based on used gasoline

²¹ based on the gasoline used and the CH ₂ content of the methanol reacted

Table 7 (Example 22 to 24) use products

H-AGO

Water H-AGO

methanol

water

2)

0.9

8.9 18.5 16.8

1.6

Σ cracking gas <C ₄ 46.4 54.3 63.6 Benzene toluene xylene 2.6 3.1 2.2 0.9 3.5 4.3 2.2 4.1 4.6 · Pyrölysebehzin <200 ⁰ C 45.7 42.1 30.9 Pyrolysis oil> 200 ° C 7.9 3.6 5.5 CO CO ₂ 0.2 0.1 5.2 0.3 4,5 0,2 Methanol (% by mass) - 80.5 86.0

H-AGO

methanol

water

Temperature ( ⁰ C) feedstock / water mass ratio catalyst load (v / vh)

- mithydr.AGO

- with methanol

- with water

680 1: 1.3

5.3 2.6

2.7 680 1: 0.6: 0.6

5.5 2.6 1.6 1.3

680 1: 0.4: 0.4

3.1 1.7 0.8 0.6

Yields (% by mass)

hydrogen

methane

ethylene

propylene

butadiene

0.3

3.9 12.1 13.9

1.4 2)

0.5

5.4 10.4 12.0

1.1

H-AGO: hydrogenated atmospheric gas oil ¹¹ based on H-AGO used

²¹ based on used H-AGO and the CH ₂ portion of the reacted

2)

0.6

7.2 14.7 16.5

1.2

Σ cracking gas <C ₄ 39.3 39.7 54.7 Benzene toluene xylene 2.8 4.5 3.5 1.7 4.2 5.3 2.4 6.8 5.5 Pyrolysis gasoline <200 ° C 33.5 33.9 31.4 Pyrolysis oil> 200 ° C 27.1 26.4 13.9 CO CO ₂ 0.2 0.1 1.9 0.1 3.7 0.1 Methanol (% by mass) - 80 91

methanol

Table 8 (Example 25 to 26)

H-AGO H-AGO

Methanol feed water water

Temperature (X) 710 710

Feedstock / water

Mass ratio 1: 1.3 1: 0.6: 0.6

catalyst loading

(v / vh) 5.3 5.5

- with hydr.AGO 2,6 2,6

With methanol - 1.6

- with water 2.7 1.3

Yields (% by mass) _ 1) 2)

Hydrogen 0.5 0.9

Methane 6.0 8.8

Ethylene 16.5, 15.8

Propylene 14.0 15.0

Butadiene 2.2 "2,4

Fission gas <C ₄ 47.1 57.9 Benzene toluene xylene 4.4 5.6 2.3 2.2 5.6 • 2.7 Pyrolysis gasoline <200 ° C 33.1 30.7 Pyrolysis oil> 200 ° C 19.8 11.4 CO CO ₂ 0.2 0.1 4.3 0.1 Methanol (% by mass) - 82

¹¹ based on H-AGO used

²¹ based on H-AGO used and the CH ₂ content of the methanol reacted

Claims (5)

- 1 - OOÖ / * f Invention claim: 1. A process for the preparation of lower olefins and aromatics, characterized in that mixtures of methanol with liquid and / or gaseous hydrocarbons in a ratio of 10: 1 to 1: 5 in the presence of a high-siliceous zeolite with channel structure at temperatures in the range of 600 to 800 ⁰ C, preferably in the range of 650 to 740 ° C, at a load of 0.5 to 6.0v / vh, at a ratio of fed steam to the feed of 0.5 to 10: 1 and at a pressure of 0.1 up to 2.0MPa are implemented. 2. The method according to item 1, characterized in that mixtures of methanol with hydrocarbons of the gasoline to gas oil boiling range are converted. 3. The method according to item 1 and 2, characterized in that instead of the methanol dimethyl ether or a mixture consisting of methanol and dimethyl ether mixture is used. 4. The method according to item 1 to 3, characterized in that a system with an effective pore diameter of 0.4 to 0.6 nm and a modulus of 5 = 35 is used as the high-silica zeolite with channel structure. 5. The method according to item 1 to 5, characterized in that in the methanol conversion in situ forming water vapor is used for the partial or total substitution of the process steam for the hydrocarbon cleavage.