CN113493699B

CN113493699B - Method for producing aromatic hydrocarbon and/or liquid fuel from light hydrocarbon

Info

Publication number: CN113493699B
Application number: CN202010265633.0A
Authority: CN
Inventors: 张爱华; 王辉; 丽萨·阮; 单军军; 乔舒亚·迈尔斯; 路易斯·吉伦; 刘华
Original assignee: China Energy Investment Corp Ltd; National Institute of Clean and Low Carbon Energy
Current assignee: China Energy Investment Corp Ltd; National Institute of Clean and Low Carbon Energy
Priority date: 2020-04-06
Filing date: 2020-04-06
Publication date: 2023-09-15
Anticipated expiration: 2040-04-06
Also published as: CN113493699A; WO2021203816A1

Abstract

The invention provides a method for producing aromatic hydrocarbon and/or liquid fuel by light hydrocarbon dehydroaromatization, which comprises the following steps: 1) Under the dehydrogenation reaction condition, carrying out dehydrogenation reaction on the light hydrocarbon stream to obtain a stream containing olefin; 2) Under the condition of aromatization reaction, the olefin-containing stream is contacted with an aromatization catalyst to carry out oligomerization/aromatization reaction, so as to obtain a stream containing aromatic hydrocarbon and or liquid fuel. The method is characterized in that dehydrogenation and oligomerization/aromatization are separated, different conditions are used for step execution, and a zeolite molecular sieve loaded with active metal components is used as an aromatization catalyst, so that the method can obtain higher alkane conversion rate, higher carbon utilization rate and higher single-cycle aromatic hydrocarbon production capacity under the preferential low-temperature aromatization condition; and the catalyst is slow in deactivation, long in one-way service life, easy to regenerate and good in multi-cycle stability.

Description

Method for producing aromatic hydrocarbon and/or liquid fuel from light hydrocarbon

Technical Field

The present invention relates to a process for producing aromatic hydrocarbons and/or liquid fuels from light hydrocarbons.

Background

Since the mid 2000 revolution in shale gas has resulted in an exponential increase in north american Natural Gas (NG) and Natural Gas Liquids (NGLs) production. This directly results in the price of NGL components, particularly ethane, reaching a low valley in recent years. The abundance and low cost of ethane also translates into the conversion of ethylene production from conventional naphtha feed to ethane feed, and the price tends to be lower as the supply increases. On the other hand, switching from catalytic reforming of naphtha to ethane cracking also results in insufficient aromatics production and increased prices. A larger drop in price with ethane and ethylene is objectively formed.

The direct conversion of light hydrocarbons to aromatic compounds has long been one of the main interests in the academia and industry. Shell Oil, exxon Mobil, SABIC and other inventors have been granted or are applying for a number of related patents for the last decades. From the point of view of the chemical reaction mechanism, the conversion of ethane to aromatics requires a step of passing through ethylene as an intermediate product. I.e. ethane, first needs to be dehydrogenated to activate to the highly reactive intermediate ethylene. Then, ethylene is oligomerized/aromatized at the acid position to an aromatic compound or the like. The dehydrogenation step is typically accomplished by the dehydrogenation function of noble metals such as Pt, pd, etc., and oligomerization/aromatization is typically accomplished over a zeolite catalyst such as ZSM-5. Thus, the direct conversion of ethane to aromatic hydrocarbon products by a one-step process needs to be accomplished in the presence of a bifunctional catalyst. However, due to thermodynamic limitations, alkane dehydrogenation needs to be accomplished at higher temperatures, such as ethane dehydrogenation, with ideal conditions being above 750 ℃. On the other hand, ZSM-5 molecular sieve catalyzed olefin aromatization is an exothermic process, with temperatures in the range of 400-500 ℃ being suitable, excessive temperatures leading to coking and the formation of substantial amounts of cleavage product methane, thereby rapidly deactivating the catalyst.

In view of this, current one-step process development is typically carried out at a temperature in the range of 550-650 ℃, which is obviously a compromise choice for compromise of the two-step reaction conditions. At this temperature, the ethane conversion is limited, with the highest equilibrium ethane to ethylene conversion being less than 30%, but for the molecular sieve catalyzed conversion between hydrocarbons, this temperature is already too high and carbon deposition and by-product formation is severe.

The main disadvantage of the one-step process in terms of product composition is the very high selectivity of the cracked product methane and heavy components (generally aromatic components with carbon numbers greater than 10), resulting in a low carbon utilization. For example, in the results disclosed in US8946107B2, methane selectivity is as high as 38%, methane selectivity can be reduced to 24% by adding 0.08% Fe, but at the same time ethane conversion is reduced by about 10%. As another example, U.S. Pat. No. 3,262B 2 discloses that the WHSV is 1.0g-C at 540-560 DEG C ₂ H ₆ The ethane aromatization results obtained from the fluidized bed reactor at/g-cat.hr were only about 30% for average ethane conversion in three cycle life runs with a total aromatics selectivity (a6+) of approximately 70% and a heavy aromatics component (a10+) of approximately 20% of the total carbon-based product, accounting for 30% of the total aromatics product. Resulting in loss of carbon-based feedstock.

Another disadvantage of the dual function catalysts used in the one-step process is that noble metals such as Pt sinter at high temperatures, and after several cycles, regeneration becomes more difficult and the catalyst activity decreases. In extreme cases, a regenerative means of chloridizing redispersion is required, but the effect is also limited.

Disclosure of Invention

The invention aims to solve the problems of low aromatic hydrocarbon yield, low carbon utilization rate, easy catalyst deactivation and the like in the process of converting light hydrocarbons into aromatic hydrocarbons in the prior art, and provides a novel method for producing aromatic hydrocarbons by dehydrogenating and aromatizing the light hydrocarbons.

As previously mentioned, temperature selection in one-step processes is a two-way problem, and the major disadvantage of current one-step processes for converting alkanes to aromatics is (1) the limited ethane conversion due to temperature limitations; (2) The catalyst is deactivated rapidly by coking, thus requiring frequent regeneration; (3) the methane content of the cracking product is higher; (4) The high proportion of heavy aromatic components such as naphthalene in the liquid product, since methane and heavy aromatic components are undesirable byproducts, their formation results in reduced carbon utilization; (5) Noble metals such as Pt as a main catalyst component are expensive, easy to sinter, and difficult to regenerate; (6) The catalyst life per pass is only a few hours and thus the use of complex reactor systems such as fluidized beds or moving beds must be considered in the process design, which will inevitably increase the complexity of the process as well as the equipment and operating costs.

The inventors of the present invention found that by separating dehydrogenation from oligomerization/aromatization, which is performed in steps, high alkane conversion and liquid aromatic yield can be easily obtained, and that noble metal catalyst is not required in the whole two-step process, so that catalyst cost can be effectively reduced.

The invention provides a method for producing aromatic hydrocarbon by dehydrogenating and aromatizing light hydrocarbon, which comprises the following steps:

1) Under the dehydrogenation reaction condition, carrying out dehydrogenation reaction on the light hydrocarbon stream to obtain a stream containing olefin;

2) Under the condition of aromatization reaction, contacting an olefin-containing stream with an aromatization catalyst to carry out oligomerization/aromatization reaction to obtain a stream containing aromatic hydrocarbon and/or liquid fuel;

wherein the aromatization catalyst comprises a zeolite molecular sieve and an active metal component supported on the zeolite molecular sieve and optionally a binder.

As described above, the present invention breaks the thought-setting of the prior art using a bifunctional catalyst by separating dehydrogenation from oligomerization/aromatization, step 1) may include conventional and well-established processes of catalytic dehydrogenation of ethane, thermal dehydrogenation, and steam cracking. Step 2) the olefins (mainly ethylene) produced from step 1) can be selectively converted to aromatic compounds, such as benzene, toluene, xylenes and/or other liquid hydrocarbons, under low temperature conditions, particularly in the range of 350-550 ℃, using a metal such as Ga-modified zeolite molecular sieve, such as ZSM-5, as a catalyst. The invention is executed step by using different conditions, and the following effects are obtained: 1) Light hydrocarbon conversion can be independently controlled by varying the conditions of the dehydrogenation zone to achieve high alkane conversion and slowest deactivation; 2) The product selectivity can be controlled by process conditions in the oligomerization/aromatization step; 3) By using the zeolite molecular sieve catalyst loaded with metal, the reaction temperature can be further reduced, and the yields of aromatic hydrocarbon and liquid oil products can be improved. Thus, the method can obtain (1) higher alkane conversion rate, (2) higher carbon utilization rate and (3) higher single-cycle aromatic hydrocarbon generation amount; and the catalyst is slow in deactivation, long in one-way service life, easy to regenerate and good in multi-cycle stability. Because of low temperature operation and long catalyst life, the fixed bed reactor can be selected for industrial production, which makes the process design simple and feasible and has low cost. In view of these technical advantages and proven catalyst performance, it is expected that the process of the present invention will be scaled up in pilot scale to achieve a smaller technical barrier in commercial scale production.

Drawings

FIG. 1 is a process flow diagram of one embodiment of the method of the present invention.

Figure 2 is a graph of the results of the BTX yield and product selectivity calculated over time based on ethane for example 1 through a series of reactors.

FIG. 3 is a graph showing the results of the one-step BTX yield over time in comparative example 1 and comparative example 2.

Figure 4 is a comparison of BTX yields over time at 450 ℃ calculated on ethane over gallium-containing versus gallium-free catalysts by series reactors.

FIG. 5 is a graph of the calculated ethylene conversion, BTX yield, methane selectivity (a) and BTX component selectivity profile (b) over time, calculated from ethylene, during the first cycle operation in a 3bar multiple cycle life experiment at 450 ℃.

Fig. 6 is a graph of the results of the catalyst cycle stability in the experiment of fig. 5, where (a) is the time required for BTX yield or ethylene conversion to drop to a certain level, (b) is the ethylene conversion capacity of the catalyst, and (c) is the BTX production capacity of the catalyst.

Detailed Description

The endpoints and any values of the ranges disclosed herein are not limited to the precise range or value, and are understood to encompass values approaching those ranges or values. For numerical ranges, one or more new numerical ranges may be found between the endpoints of each range, between the endpoint of each range and the individual point value, and between the individual point value, in combination with each other, and are to be considered as specifically disclosed herein.

Fig. 1 is a simplified process flow diagram showing the method of the present invention. As shown in fig. 1, light hydrocarbons (light alkanes, having boiling points not exceeding-10 ℃) such as ethane will first be dehydrogenated to produce the corresponding olefins such as ethylene, which will then be oligomerized/aromatized to produce aromatics and other liquid fuels.

With respect to this two-step process, the inventors of the present invention have mentioned in the prior patent literature that the conversion of lower alkanes to aromatics is split into two steps, with the objective of step 1) alkane dehydrogenation to achieve suitably high olefin concentrations in step 2) feed and the objective of step 2) to achieve high aromatics and other liquid fuel production and system efficiency by catalytic conversion of the olefin-containing feed.

The present invention will continue and extend the two-step process concept, focusing on step 2) operating at lower temperature conditions, with the aim of further improvements in BTX yield, catalyst life, overall carbon utilization and catalyst long term stability, optimizing process conditions and catalyst performance. In addition, the product composition is not limited to aromatic compounds, and may include hydrocarbons in the gasoline range. The process can be designed for different modes of operation, for example, with the aim of maximizing aromatic compounds, with the aim of maximizing gasoline, and with the aim of optimizing overall catalyst performance and system efficiency. The choice of operating mode may depend on market, demand and economics.

By separating the ethylene aromatization (step 2) from the ethane dehydrogenation (step 1), both steps can be operated under their optimal conditions.

The invention is thatIn step 2) it is preferred to operate at lower temperature conditions, preferably 300-600 ℃, further preferably 350-600 ℃, more preferably 400-550 ℃. The inventors of the present invention found that higher temperatures favor the formation of BTX, but at the same time CH ₄ And the formation of heavy aromatic hydrocarbon products (a10+) having a carbon number greater than 10, resulting in a decrease in carbon utilization, defined and calculated as the sum of the amounts of light aromatic hydrocarbon (A6-A9) and lower hydrocarbon (C2-C5) products.

Specifically, when BTX is the target product, the temperature in step 2) is preferably 450 to 550 ℃.

When other liquid fuels such as gasoline are used as target products, the temperature in step 2) is preferably 350-450 ℃.

In the present invention, liquid fuel means hydrocarbon component of C5 or more, i.e., gasoline component.

The temperature of step 2) is preferably 400-500 c, with the aim of maximizing the overall catalyst performance and system efficiency.

Step 2) may be carried out under normal pressure or at a pressure of 1 to 5 bar.

The liquid hourly space velocity (WHSV) of the olefin-containing stream in terms of ethylene may be in the range of from 0.5 to 10g/g-cat/hr, preferably from 0.75 to 3g/g-cat/hr.

Step 2) may be carried out in a conventional fixed bed reactor, fluidized bed reactor, moving bed reactor.

The catalyst of step 2) is a zeolite molecular sieve with or without an active metal component and a binder for catalyst shaping.

The active metal component may be one or more of the group IIIA, VIII, VIB metal elements of the periodic Table of the elements, for example Ga, fe, ni, ag and Mo, preferably one or more of Ga, ni. The active metal component is present in an amount of from 0.4 to 5% by weight, preferably from 0.8 to 2.5% by weight, based on the total amount of catalyst, for example 0.4, 0.5, 0.8, 1, 1.2, 1.4, 2, 2.1, 2.2, 2.3, 2.4, 2.5% by weight, based on the element.

In the present invention, the zeolite molecular sieve may be various molecular sieves having a zeolite structure, preferably, the zeolite molecular sieve has a silica alumina molar ratio of 5 to 300, and more preferably, one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, Y-type zeolite, beta zeolite, ferrierite (Ferrierite) and mordenite.

The zeolite molecular sieve is preferably present in an amount of not less than 50 wt%, more preferably above 70 wt%, for example 75-85 wt%, based on the total amount of catalyst.

The aromatization catalyst may be a powder or a molded body, preferably the aromatization catalyst is a molded body containing an active metal, a zeolite molecular sieve and a binder. The binder may be any of a variety of materials capable of shaping the zeolite molecular sieve into a shaped catalyst, and may be, for example, one or more of silica, alumina, clay, aluminum phosphate, zirconia. Preferably, the weight ratio of zeolite molecular sieve to binder is from 50 to 99:1 to 50, preferably from 70 to 90:10 to 30, more preferably from 75 to 85:15 to 25.

According to a preferred embodiment of the present invention, the active metal component is present in an amount of 0.4 to 4 wt.%, preferably 0.8 to 2.4 wt.%, the zeolite molecular sieve is present in an amount of 50 to 99 wt.%, preferably 70 to 90 wt.%, and the binder is present in an amount of 0.6 to 49.6 wt.%, preferably 7.6 to 29.2 wt.%, based on the total amount of aromatization catalyst.

The catalyst may be commercially available or may be prepared by a known method. For example, the active metal component is first supported on the zeolite molecular sieve by an impregnation method, such as a saturation impregnation method, and then the zeolite molecular sieve is extruded with a binder by a conventional extrusion method, for example, into a bar, column or clover shape.

According to a preferred embodiment of the present invention, the synthesis of the aromatization catalyst is carried out by wet impregnation of zeolite molecular sieve powder with a binder such as boehmite, water and a nitrate solution of the active metal component to make a paste. The paste is extruded, dried and calcined at 500-650 ℃ for 8-12 hours. Crushing and sieving the calcined extruded strips, and taking particles with 20-40 meshes for catalytic reaction.

In the present invention, the dehydrogenation reaction of step 1) may be any one of (ethane) thermal dehydrogenation, steam cracking, catalytic dehydrogenation and Oxidative Dehydrogenation (ODH).

Since the dehydrogenation reaction is carried out separately, the conditions of the dehydrogenation reaction can be set such that a relatively high conversion (e.g., > 50%) can be achieved and the temperature can be significantly lower than the typical ethane cracker outlet temperature (typically >850 ℃), the dehydrogenation temperature being, for example, about 750 ℃. Because of the relatively low temperature, coking is negligible in the absence of steam in this zone, and thus no steam may be vented. In order to facilitate reducing the partial pressure of the alkane to facilitate conversion of the light hydrocarbons, a diluent, such as nitrogen, is preferably introduced into the reactor.

If no steam is used in the dehydrogenation, the product stream can be used directly in the next oligomerization/aromatization reaction without cooling separation to remove moisture and reheating to effect the reaction. This will save the cost of the overall process. Since the oligomerization/aromatization reactions are also carried out independently, the product stream from the dehydrogenation reaction (typically an alkane/alkene/hydrogen mixture) can be selectively converted to an aromatic compound or gasoline product by selecting the appropriate catalyst and controlling the reaction conditions. Since only oligomerization/aromatization reaction and low-carbon alkane dehydrogenation reaction requiring high activation energy are needed in the step, the catalyst used can not contain noble metals with strong dehydrogenation function such as Pt, and further can reduce main side reactions such as hydrogenation of ethylene into ethane and further hydrogenolysis into methane, and the like, so that the generation amount of methane is obviously reduced.

In the present invention, the dehydrogenation reaction may be performed in the presence of a dehydrogenation catalyst or may be performed in the absence of a dehydrogenation catalyst.

The dehydrogenation catalyst may be various catalysts having an alkane dehydrogenation function, and preferably, the dehydrogenation catalyst is a supported catalyst comprising a support and a metal component having dehydrogenation activity supported on the support, the support being an inorganic refractory oxide having no acid center.

Preferably, the metal component having dehydrogenation activity is a noble metal component such as Pt or Pd.

According to one embodiment of the present invention, the content of the metal component having dehydrogenation activity is 0.01 to 2.0 wt%, preferably 0.02 to 0.2 wt%, based on the total amount of the dehydrogenation catalyst.

Preferably, the carrier is one or more of silica, alumina, silicon carbide, clay, cerium oxide, lanthanum oxide, magnesium oxide, titanium oxide and zirconium oxide.

The catalyst may be commercially available or may be prepared by a known method.

As described above, since the dehydrogenation reaction is independently performed in the present invention, the dehydrogenation reaction may be performed under conditions favorable for dehydrogenation, and preferably, the dehydrogenation reaction is performed in the presence of a dehydrogenation catalyst at a temperature of less than 900 ℃, preferably 650 to 850 ℃, and more preferably 650 to 800 ℃. This temperature is lower than the typical ethane cracking reaction temperature (typically greater than 850 ℃).

Preferably, the light hydrocarbon stream has a GHSV of 500-20000h ^-1 Preferably 800-5000h ^-1 。

According to another embodiment of the invention, the dehydrogenation reaction is carried out in the absence of the dehydrogenation catalyst. At this time, the temperature of the dehydrogenation reaction is preferably 700 to 900 ℃. The residence time of the reactants is preferably from 0.05 to 30 seconds. In the present invention, the residence time of the reactants means the residence time of the reactants at the above-mentioned dehydrogenation reaction temperature of 700 to 900 ℃, i.e., the dehydrogenation reaction time.

In order to facilitate reducing the partial pressure of the alkane to facilitate conversion of the light hydrocarbon, preferably a diluent is introduced into the dehydrogenation reactor. The diluent may be, for example, an inert gas such as nitrogen which does not adversely affect the reaction.

Since step 1) of the present invention is to dehydrogenate light hydrocarbons with the objective product being the corresponding olefin, the objective of the dehydrogenation reaction is to obtain as much olefin as possible, and the product of the dehydrogenation reaction is referred to as "olefin-containing stream". The olefin-containing stream may be subjected directly to the oligomerization/aromatization reaction of step 2) without separation, thereby saving time for cooling the separation and reheating to the temperature required for the oligomerization/aromatization reaction, shortening the flow, and greatly reducing the costs incurred thereby. Thus, preferably, the olefin-containing stream is directly subjected to the oligomerization/aromatization reaction of step 2) without separation.

In the present invention, the oligomerization/aromatization reaction in step 2) refers to the reaction in which the olefin stream obtained by dehydrogenation in step 1) undergoes oligomerization and aromatization to form aromatic hydrocarbons. The raw materials for the oligomerization/aromatization reaction in the step 2) may be all derived from the step 1), or may be obtained by adjusting the raw materials based on the product of the step 1) as needed, preferably so that the ethylene content is not less than 20% by volume, preferably 20 to 50% by volume.

According to the present invention, the hydrocarbons in the light hydrocarbon stream may be various substances capable of undergoing dehydrogenation to undergo oligomerization/aromatization to produce aromatic hydrocarbons, for example, various alkanes having no more than 5 carbon atoms. Preferably, the ethane content in the light hydrocarbon stream is not less than 65% by volume, more preferably 75 to 100% by volume.

Preferably, the ethylene content of the predominantly olefin-containing stream is not less than 20% by volume, preferably from 20 to 50% by volume.

Preferably, the dehydrogenation reaction of step 1) and the oligomerization/aromatization reaction of step 2) are carried out in different zones of the same reactor or in different reactors. In order to enable the reaction to proceed continuously without interruption of the operation due to deactivation of the catalyst, it is preferred that at least two of the dehydrogenation reactions described in step 1) and the oligomerization/aromatization reactions described in step 2) are each arranged in parallel, whereby the reaction can be switched into another reactor after deactivation of the catalyst, the original reactor being subjected to regeneration of the deactivated catalyst. The reactor may be a fixed bed reactor or a fluidized bed reactor.

In order to obtain a liquid fuel such as aromatic hydrocarbon, gasoline product with higher purity and to convert light hydrocarbon into target product as much as possible, the method according to a preferred embodiment of the present invention preferably further comprises subjecting the liquid fuel-containing stream obtained in step 2) to gas-liquid separation as shown in fig. 1 to obtain a liquid fuel-containing liquid phase stream which is discharged as a product or fed to a subsequent separation process.

Preferably, the separated gas phase stream is further subjected to gas separation to obtain a hydrogen stream, a fuel stream and a light hydrocarbon stream.

It is further preferred that the light hydrocarbon stream (C2-C4 component) is returned to step 1) as a feedstock for the dehydrogenation reaction.

The present invention will be described in detail by examples. In the following examples, the molecular sieves used in the oligomerization/aromatization catalysts are commercially available from Zeolyst corporation. The active metal component is present in a weight percent relative to the molecular sieve unless otherwise indicated.

In the following calculation of the results, there are two kinds of the feed amount based on ethane (also referred to as ethane feed amount) and the feed amount of ethylene (also referred to as ethylene feed amount). All "amounts" are defined as molar amounts of carbon.

The ethane conversion was calculated as: ethane conversion,% = 100% × (ethane feed-post-reaction ethane feed)/ethane feed

The calculation formula of the ethylene conversion rate is as follows: ethylene conversion,% = 100% × (ethylene feed-ethylene feed after reaction)/ethylene feed

The calculation formula of the component selectivity is as follows: selectivity of component Y = 100% × (amount of component Y produced/amount of ethane or ethylene reacted)

The yield of component Y was calculated as: conversion of ethane or ethylene x selectivity of component Y (based on ethane or ethylene)

In order to calculate the ethylene conversion capacity and BTX production capacity, several lower calculation limits (somewhat "lower limits") are typically defined in terms of ethylene conversion or BTX yield in the resulting curves of ethylene conversion and BTX yield over time. For example, defining a lower BTX yield limit of 30% calculated based on ethylene, the time from the start of the reaction to the "lower limit", i.e. the single pass catalyst life (a bit of "life"), can be obtained in a life experiment and the average conversion over this reaction period calculated.

The ethylene conversion capacity was calculated as: lifetime up to a defined lower limit x average conversion calculated over this reaction time based on ethane or ethylene x liquid hourly space velocity of ethane or ethylene feed

The BTX generation capacity is calculated as: lifetime up to a defined lower limit x average BTX yield calculated over the ethane or ethylene over this reaction time x liquid hourly space velocity of ethane or ethylene feed

The composition of the product was measured on-line or off-line using gas chromatography.

Example 1

Using the process shown in fig. 1, a volume ratio of C of 1.67:1 was used ₂ H ₆ /N ₂ The mixed gas is sent into two reactors of dehydrogenation and oligomerization/aromatization in series, the dehydrogenation temperature is set to 750 ℃, and the GHSV of ethane is 1000h ^-1 The pressure is normal pressure, and the reaction is carried out under the condition of no catalyst. Under these conditions, the conversion of ethane was about 57% and the concentration of ethylene in the dehydrogenation product gas stream was about 22% by volume. The oligomerization/aromatization fixed bed reactor set-up temperatures were 400 ℃, 450 ℃, 500 ℃ and 550 ℃, respectively, whsv=2.68 g-C ₂ H ₆ /g-cath hr (corresponding to 1.6g-C ₂ H ₄ Per g-cat hr) under normal pressure (1 bar), and the catalyst is Ga/ZSM-5/Al prepared by saturated impregnation method ₂ O ₃ Catalyst (Ga content of 2.0 wt%, ZSM-5 silicon-aluminum mol ratio of 30, molecular sieve and binder Al) ₂ O ₃ Is 82.5/17.5). The BTX yield and product selectivity results calculated based on ethane are shown in table 1 and figure 2.

C shown in FIG. 2 (a) ₂ H ₄ Selectivity represents unconverted C in the oligomerization/aromatization reactor ₂ H ₄ Is a combination of the amounts of (a) and (b). The data show that as the temperature increases, the amount of unconverted ethylene increases, indicating rapid catalyst deactivation at high temperature conditions. C under four temperature conditions ₂ H ₄ The run times for the selectivity increase to 20% were respectively: 5190 minutes at 400 ℃, 4291 minutes at 450 ℃, 2530 minutes at 500 ℃, and 1159 minutes at 550 ℃.

As can be seen from fig. 2 (b), the yield of BTX is also greatly affected by temperature. The higher the temperature, the higher the initial yield of BTX, but also the faster the decrease due to coking. The run times at which BTX yield dropped to the lower limit of 15% at these four temperatures were respectively: 2705 minutes at 400 ℃, 3772 minutes at 450 ℃, 2747 minutes at 500 ℃ and 1494 minutes at 550 ℃.

In FIG. 2 (c) is shown CH under four temperature conditions ₄ Selectivity. CH measured at 400 DEG C ₄ The selectivity was about 8%. Due to the series mode operation, the 8% CH ₄ Most of the results from step 1, the added portion is newly formed in step 2 oligomerization/aromatization. Corresponding to BTX production, the initial selectivity for methane is higher at high temperature, but also decreases rapidly with catalyst deactivation.

In the BTX product, the benzene fraction is greatly affected by temperature, as shown in fig. 2 (d), the higher the temperature, the higher the B/BTX. Regarding the value of benzene in the product, it depends on the definition of the target product. Benzene has higher value if light aromatic hydrocarbon BTX is used as a target product. However, when gasoline is used as the target product, the benzene content is strictly limited, preferably at a low value.

The BTX production capacity with 15% BTX yield as lower limit was 22.5 at 400 ℃, 40.3 at 450 ℃, 34.0 at 500 ℃ and 18.4g-BTX/g-cat at 550 ℃ calculated from the curve of BTX yield over time (fig. 2-b), see table 1.

In summary, high temperature conditions can promote more BTX production at the expense of a short catalyst life and more methane formation of the cleavage product.

Comparative example 1

The conversion of ethane was carried out as in example 1, except that the one-step ethane aromatization was carried out using a Pt/ZSM-5 catalyst. Specifically, a mixed gas of ethane and nitrogen is used as a feed. The dehydrogenation/oligomerization/aromatization temperature was set at 630 ℃, whsv=1.34 g-C ₂ H ₆ Per g-cat.hr at atmospheric pressure, the catalyst was Pt/ZSM-5/Al with a Pt loading of 0.05 wt% ₂ O ₃ (ZSM-5 has a silica-alumina molar ratio of 30, molecular sieve and binder Al) ₂ O ₃ 70/30 by weight). The results are shown in table 1 and (a) in fig. 3. As can be seen from FIG. 3 (a), although the WHSV is only half of example 1, the BTX yield decreases rapidly with the reaction time, oneThe single cycle operation can only last about 400 minutes where the BTX yield has fallen to the lower limit of 15%.

Comparative example 2

Ethane conversion was performed as in comparative example 1, except that one-step ethane aromatization was performed using a Pt-Zn-Sn/ZSM-5 catalyst. Specifically, a mixed gas of ethane and nitrogen is used as a feed. The dehydrogenation/oligomerization/aromatization temperature was set at 630 ℃, whsv=1.34 g-C ₂ H ₆ Per g-cat.hr at normal pressure, catalyst Pt loading of 0.05 wt%, zn loading of 0.017 wt%, sn loading of 0.03 wt% Pt-Zn-Sn/ZSM-5/Al ₂ O ₃ (ZSM-5 has a silica-alumina molar ratio of 30, molecular sieve and binder Al) ₂ O ₃ 70/30 by weight). The results are shown in Table 1 and FIG. 3 (b). As can be seen from FIG. 3 (b), the BTX yield decreased more slowly than in FIG. 3 (a), but the reaction could also be run for only 400 minutes.

Furthermore, comparing the single cycle BTX production capacities of both the example 1 and comparative examples 1,2, it was found that the value of the 1-step process (comparative examples 1, 2) was not higher than 5g-BTX/g-cat; whereas by the two-step process, the single cycle BTX production capacity was 18-40g-BTX/g-cat over the temperature range of 400-550℃and several times that of the one-step process, see Table 1.

Comparative example 3

Ethane conversion was performed as in example 1, except that the ethylene aromatization catalyst used was ZSM-5 without metal addition (ZSM-5 was the same as in example 1, molecular sieves and binder Al) ₂ O ₃ 82.5/17.5) by weight, the reaction being carried out at a temperature of 450 c only. The results of the BTX yield over time are shown in fig. 4. The BTX production capacity with 15% BTX yield as the lower limit was 16.5g-BTX/g-cat calculated from the curve.

TABLE 1

Example 2

Unlike the process of example 1, the product according to step 1) of example 1 is usedA mixture gas flow with simplified formulation and a composition and volume ratio of C ₂ H ₄ /H ₂ /N ₂ =1:1:1. The catalyst is the same as the oligomerization/aromatization catalyst of example 1, namely Ga/ZSM-5/Al prepared by the saturation impregnation method ₂ O ₃ Catalyst (Ga content of 2.0 wt%, ZSM-5 silicon-aluminum mol ratio of 30, molecular sieve and binder Al) ₂ O ₃ Is 82.5/17.5). The aromatization conditions were 450 ℃ and 550 ℃, normal pressure, whsv=0.75 g-C ₂ H ₄ The reaction time was 6 hours per g-cath. The results of the 6 hour average BTX yield and product selectivity calculated based on ethylene are shown in table 2.

Comparative example 4

Experiments were performed as in example 2, except that the catalyst used for aromatization of ethylene was ZSM-5 with no metal addition (ZSM-5 was the same as in example 1). The average results of the reaction at 450℃for 6 hours are shown in Table 2.

Example 3

An experiment was performed as in example 2, except that the catalyst used for aromatization of ethylene was ZSM-5 containing 1% nickel by weight based on the total weight of the molecular sieve (ZSM-5 was the same as in example 1). The average results of the reaction at 450℃for 6 hours are shown in Table 2.

Example 4

An experiment was performed as in example 2, except that the catalyst used for aromatization of ethylene was ZSM-5 containing 1% iron by weight based on the total weight of the molecular sieve (ZSM-5 was as in example 1). The average results of the reaction at 450℃for 6 hours are shown in Table 2.

Example 5

An experiment was performed as in example 2, except that the catalyst used for aromatization of ethylene was ZSM-5 containing 1% silver by weight based on the total weight of the molecular sieve (ZSM-5 was the same as in example 1). The average results of the reaction at 450℃for 6 hours are shown in Table 2.

TABLE 2

As can be seen from the data of example 2 in table 2 above, BTX yields can be as high as 63.5% at 550 ℃; CH (CH) ₄ And the selectivity to heavy aromatics (a10+) was 11.7% and 9.0%, respectively. The collective selectivity of C2-C5 and A6-A9 can be calculated by addition, and the carbon utilization is about 80% under the reaction conditions according to the carbon utilization defined hereinabove, i.e. (C2-C5) + (A6-A9).

At 450℃the BTX yield was 55.1% lower than 550℃by 8.4%, but CH ₄ And A10+ selectivity is also lower, carbon utilization is about 90%, about 10% higher than 550 ℃.

Since the components in C2-C5 are predominantly light hydrocarbons ethane, propane, these components can be returned to the dehydrogenation reactor via the circulation path shown in fig. 1 to produce more liquid product, high carbon utilization is also a factor to be considered in optimizing process conditions.

In the liquid product, the benzene fraction is significantly higher than the value of 450 ℃ at 550 ℃, so if high benzene content is targeted, the reactor needs to be operated at high temperature conditions.

As can be seen from the comparison of table 2 and fig. 4, in the catalytic reaction without gallium addition (comparative example 3, comparative example 4), the BTX yield was lower than that of example 1 and example 2, respectively, in combination with fig. 2, while the selectivity of propane, butane, etc. was significantly higher than that of the former (i.e., example 2) in the product, thereby revealing that the main efficacy of gallium was to increase the BTX yield by aromatizing propane, butane.

Example 6

An experiment was performed as in example 2, except that the conditions for aromatization of ethylene were 450 ℃, 3bar, whsv=1.5 g-C ₂ H ₄ C/g-cat hr, volume ratio of 0.67:1:1:1 ₂ H ₆ /C ₂ H ₄ /H ₂ /N ₂ Is a gas mixture of (a) and (b). The reaction was subjected to a multi-cycle life test of 3 months. Each cycle operation included about 100 hours of reactionAnd about 24 hours of catalyst regeneration. The catalyst regeneration used a 50:50 volume ratio air/nitrogen mixture.

For each cycle, catalyst life, ethylene conversion capacity and BTX production capacity were calculated separately, with the lower limits defined as (a) 30% BTX yield calculated based on ethylene and (b) 80% ethylene conversion.

In this long-term stability test, 17 cycles of the test were completed, cotransformed 1990g-C ₂ H ₄ Per g-cat (lower limit at 80% ethylene conversion), 858g of BTX product was produced. On average, per cycle, the ethylene conversion capacity was 117g-C ₂ H ₄ Each g-cat had a BTX production capacity of 50g-BTX/g-cat.

Fig. 5 discloses the variation over time of ethylene conversion, BTX yield, methane selectivity (a) and selectivity (b) of the various components in BTX in the first catalytic cycle. In the lifetime experiment of nearly 100 hours completed, the ethylene conversion capacity and BTX production capacity were calculated to be 127g-C, respectively, based on the two defined lower limits, namely, 30% BTX yield (lower limit 1) and 80% ethylene conversion (lower limit 2) ₂ H ₄ Per g-cat (lower limit 1) and 140g-C ₂ H ₄ Per g-cat (lower limit 1), 62g-BTX/g-cat (lower limit 2) and 66g-BTX/g-cat (lower limit 2).

Figure 6 integrates catalyst life, ethylene conversion capacity and BTX production capacity as a function of cycle number. As can be seen from fig. 6, a decrease in capacity value was observed during the first 5-6 cycles, but then stabilized. Overall, the catalyst has better long-term stability.

The above examples fully illustrate that the two-step process provided by the present invention provides significant technical advantages since the two reaction steps can be operated under individually optimized conditions. In particular, step 1) may be an ethane cracking process already established in industry, step 2) may be operated at a lower temperature range, for example 350-550 ℃, and suitable process conditions may be selected according to the target product: (1) With BTX, in particular benzene, as target product, step 2) can be carried out at a temperature in the range 500-550 ℃; (2) Taking a gasoline product as a target product, wherein the optimal temperature range of the step 2) is 350-400 ℃; (3) With the aim of BTX yield, catalyst lifetime and carbon utilization being highest, step 2) is performed at about 450 ℃ to obtain optimal performance of the overall catalyst and efficiency.

In addition, low temperature operation has other benefits, from an engineering standpoint, lower temperature operation and relatively longer catalyst life would allow for the use of simple and easy fixed bed reactors with relatively low requirements on reactor materials. All of these advantages will help reduce equipment costs (reactor construction and materials) and operating costs.

The preferred embodiments of the present invention have been described in detail above, but the present invention is not limited thereto. Within the scope of the technical idea of the invention, a number of simple variants of the technical solution of the invention are possible, including combinations of the individual technical features in any other suitable way, which simple variants and combinations should likewise be regarded as being disclosed by the invention, all falling within the scope of protection of the invention.

Claims

1. A process for producing aromatic hydrocarbons and/or liquid fuels from light hydrocarbons, the process comprising the steps of:

1) Under the dehydrogenation reaction condition, carrying out dehydrogenation reaction on the light hydrocarbon stream rich in ethane to obtain a stream containing olefin;

wherein the aromatization catalyst comprises a zeolite molecular sieve and an active metal component supported on the zeolite molecular sieve and a binder;

wherein the olefin-containing stream is directly subjected to the oligomerization/aromatization reaction of step 2) without separation;

wherein the ethane-enriched light hydrocarbon stream has an ethane content of from 75 to 100 wt.%; the ethylene content of the olefin-containing stream is from 20 to 50 wt.%;

step 2) taking BTX as a target product, wherein the temperature of the aromatization reaction is 450-550 ℃; or 2) taking liquid fuel as a target product, wherein the temperature of the aromatization reaction is 350-450 ℃;

wherein the content of the active metal component is 0.4-5 wt% based on the total amount of the aromatization catalyst, and the weight ratio of the zeolite molecular sieve to the binder is 70-90:10-30.

2. The process according to claim 1, wherein the active metal component is present in an amount of 0.8 to 2.5 wt%, based on the total amount of aromatization catalyst.

3. The method of claim 1, wherein the weight ratio of the zeolite molecular sieve to the binder is 75-85:15-25.

4. The method of claim 1, wherein the active metal component is one or more of element Ga, fe, ni, ag, mo.

5. The process of any one of claims 1-4, wherein the zeolite molecular sieve has a silica to alumina molar ratio of from 5 to 300.

6. The process of claim 5, wherein the zeolite molecular sieve is selected from one or more of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, zeolite Y, beta zeolite, ferrierite, and mordenite.

7. The process of any of claims 1-4, wherein the aromatization reaction conditions comprise: the olefin-containing stream has a WHSV, calculated as ethylene, of from 0.5 to 10g/g-cat/hr.

8. The process of claim 7, wherein the aromatization reaction conditions comprise: the olefin-containing stream has a WHSV of from 0.75 to 3g/g-cat/hr, calculated as ethylene.

9. The process of any of claims 1-4, wherein the dehydrogenation reaction of step 1) and the oligomerization/aromatization reaction of step 2) are performed in different zones of the same reactor or in different reactors.

10. The process according to any one of claims 1 to 4, wherein the process further comprises subjecting the aromatic hydrocarbon and/or liquid fuel containing stream obtained from the aromatization product of step 2) to gas-liquid separation to obtain a gas phase stream and a liquid phase stream containing aromatic hydrocarbon and/or liquid fuel.

11. The process of claim 10, wherein the separated vapor phase stream is further subjected to gas separation to provide a hydrogen stream, a fuel stream, and a light hydrocarbon stream.

12. The process of claim 11, wherein the dehydrogenation reaction is performed by returning the light hydrocarbon stream to step 1).

13. The process according to any one of claims 1 to 4, wherein the dehydrogenation reaction is carried out in the presence of a dehydrogenation catalyst which is a supported catalyst comprising a carrier and a metal component having dehydrogenation activity supported on the carrier, the carrier being an inorganic refractory oxide having no acid center.

14. The method of claim 13, wherein the support is one or more of silica, alumina, silicon carbide, clay, ceria, lanthana, magnesia, titania, zirconia; the metal component with dehydrogenation activity is Pt and/or Pd; the content of the metal component having dehydrogenation activity is 0.01 to 2% by weight based on the total amount of the dehydrogenation catalyst.

15. The process according to claim 14, wherein the content of the metal component having dehydrogenation activity is 0.02 to 0.2% by weight based on the total amount of the dehydrogenation catalyst.

16. The method of any of claims 1-4, wherein the dehydrogenation reaction conditions comprise: the temperature is 650-850 ℃, and the GHSV of the light hydrocarbon stream is 500-20000h ^-1 。

17. The method of claim 16, wherein the dehydrogenation reaction conditions comprise: the GHSV of the light hydrocarbon stream is 800-5000h ^-1 。

18. The process of any one of claims 1-4, wherein the dehydrogenation reaction is performed in the absence of a catalyst.

19. The method of claim 18, wherein the dehydrogenation reaction conditions comprise a temperature of 700-900 ℃ and a reactant residence time of 0.05-30 seconds.