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

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 hydrocarbons and/or liquid fuels from light hydrocarbons

Technical field

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

Background technique

Since the mid-2000s, the shale gas revolution has led to exponential growth in North American natural gas (NG) and natural gas liquids (NGLs) production. This has directly caused the price of NGL components, especially ethane, to reach a low point in recent years. Abundant and low-price ethane has also led to the conversion of ethylene production from traditional naphtha as the raw material to ethane as the main raw material, and due to the increasing supply, the price has also tended to be low. On the other hand, the switch from naphtha catalytic reforming to ethane cracking has also resulted in insufficient production of aromatics and increased prices. Objectively, there is a large price gap with ethane and ethylene.

The direct conversion of light hydrocarbons into aromatic compounds has long been one of the major interests in academia and industry. Over the past few decades, Shell Oil, Exxon Mobil, SABIC and other inventors have been granted or are applying for multiple related patents. From the perspective of chemical reaction mechanism, the conversion of ethane to aromatics requires the step of ethylene as an intermediate product. That is, ethane first needs to be dehydrogenated and activated to the highly reactive intermediate ethylene. Then, ethylene is oligomerized/aromatized at the acidic site and converted into aromatic compounds and the like. The dehydrogenation step usually requires the dehydrogenation function of noble metals such as Pt and Pd, while oligomerization/aromatization is usually achieved on zeolite catalysts such as ZSM-5. Therefore, direct conversion of ethane into aromatic products via a one-step process requires the presence of a bifunctional catalyst. However, due to thermodynamic limitations, alkane dehydrogenation needs to be completed at a higher temperature, such as ethane dehydrogenation. The ideal condition is above 750°C. On the other hand, the olefin aromatization reaction catalyzed by ZSM-5 molecular sieve is an exothermic process. The appropriate temperature range is 400-500°C. Excessive temperature will lead to coking and the generation of a large amount of cracking product methane, thereby causing the catalyst to rapidly Deactivated.

In view of this, the current one-step process development is generally carried out in the temperature range of 550-650°C. Obviously, this is a compromise choice to take into account the two-step reaction conditions. Under this temperature condition, the ethane conversion rate is limited, and the maximum equilibrium conversion rate of ethane to ethylene is less than 30%. However, for the conversion between hydrocarbons catalyzed by molecular sieves, this temperature is already too high, causing carbon deposition and By-product formation is severe.

In terms of product composition, the main disadvantage of the one-step method is that the selectivity of the cracked product methane and heavy components (generally refers to aromatic hydrocarbon components with a carbon number greater than 10) is very high, resulting in low carbon utilization. For example, in the results disclosed in US8946107B2, the methane selectivity is as high as 38%. By adding 0.08% Fe, the methane selectivity can be reduced to 24%, but at the same time, the ethane conversion rate is also reduced by about 10%. As another example, US10087124B2 discloses the ethane aromatization results obtained from a fluidized bed reactor at 540-560°C and a WHSV of 1.0gC ₂ H ₆ /g-cat.hr, averaged over three cycle life runs. The ethane conversion rate is only about 30%, the total aromatics selectivity (A6+) is close to 70%, and the heavy aromatic component (A10+) is close to 20% of the total carbon-based products and accounts for 30% of the total aromatic products. This results in the loss of carbon-based raw materials.

Another disadvantage of the bifunctional catalyst used in the one-step method is that precious metals such as Pt will sinter at high temperatures. After several cycles, regeneration will become increasingly difficult and the catalyst activity will decrease. In extreme cases, the regeneration method of chlorination and redispersion needs to be used, but the effect is also limited.

Contents of the invention

The purpose of the present invention is to overcome the problems in the prior art that exist in the process of converting light hydrocarbons into aromatic hydrocarbons, such as low aromatic hydrocarbon yield, low carbon utilization rate, and easy catalyst deactivation, and provide a new light hydrocarbon dehydrogenation aromatization method. The method of producing aromatic hydrocarbons can achieve higher aromatic hydrocarbon yield and carbon utilization, and can effectively solve problems such as catalyst deactivation and greatly extend the service life of the catalyst.

As mentioned before, the temperature selection of the one-step method is a dilemma. The main disadvantages of the current one-step method for converting alkanes to aromatics are (1) limited ethane conversion rate due to temperature limitations; (2) ethane conversion caused by coking The catalyst deactivates quickly and therefore requires frequent regeneration; (3) the methane content of the cracked product is high; (4) the proportion of heavy aromatic hydrocarbon components such as naphthalene in the liquid product is high, since both methane and heavy aromatic hydrocarbon components are undesirable By-products, their generation leads to a reduction in carbon utilization; (5) Precious metals such as Pt, as the main catalyst component, are expensive, easy to sinter, and difficult to regenerate; (6) The single-pass life of the catalyst is only a few hours, so in process design The use of complex reactor systems such as fluidized beds or moving beds must be considered, which will inevitably increase the complexity of the process as well as equipment and operating costs.

The inventors of the present invention have found that by separating dehydrogenation from oligomerization/aromatization and performing it in steps, high alkane conversion rates and liquid aromatic hydrocarbon yields can be easily obtained without using Precious metal catalysts can effectively reduce catalyst costs.

The invention provides a method for producing aromatic hydrocarbons by dehydrogenating aromatization of light hydrocarbons, which method includes the following steps:

1) Under dehydrogenation reaction conditions, dehydrogenate the light hydrocarbon stream to obtain an olefin-containing stream;

2) Under aromatization reaction conditions, contact the olefin-containing stream with an aromatization catalyst to perform an oligomerization/aromatization reaction to obtain a stream containing aromatic hydrocarbons and/or liquid fuel;

Wherein, the aromatization catalyst contains a zeolite molecular sieve, an active metal component supported on the zeolite molecular sieve, and an optional binder.

As mentioned above, the present invention breaks the conventional thinking of using dual-functional catalysts. By separating dehydrogenation and oligomerization/aromatization, step 1) can include catalytic dehydrogenation of ethane, thermal dehydrogenation and water Conventional and mature processes such as steam cracking. Step 2) Using a metal such as Ga-modified zeolite molecular sieve such as ZSM-5 as a catalyst, the olefins (mainly olefins) produced from step 1) can be selectively converted under low temperature conditions, especially in the range of 350-550°C. Ethylene) is converted into aromatic compounds such as benzene, toluene, xylene and/or other liquid hydrocarbons. The present invention is executed step by step using different conditions and achieves the following effects: 1) light hydrocarbon conversion can be independently controlled by changing the conditions in the dehydrogenation zone to achieve high alkane conversion rate and the slowest deactivation; 2) product selection The properties can be controlled by process conditions in the oligomerization/aromatization step; 3) By using metal-loaded zeolite molecular sieve catalysts, the reaction temperature can be further reduced and the yield of aromatics and liquid oils can be increased. Therefore, this method can achieve (1) higher alkane conversion rate, (2) higher carbon utilization rate, (3) higher single-cycle aromatic hydrocarbon production; and the catalyst deactivation is slow, the single-pass life is long, and it is easy to regenerate. , and has good multi-cycle stability. Due to low-temperature operation and long catalyst life, fixed-bed reactors can be selected for industrial production, which will make process design simple and cost-effective. In view of these technical advantages and proven catalyst performance, it is foreseeable that the method of the present invention has smaller technical barriers in the process of pilot scale-up and commercial-scale production.

Description of the drawings

Figure 1 is a process flow diagram of an embodiment of the method of the present invention.

Figure 2 is a graph showing the changes in BTX yield and product selectivity over time calculated based on ethane in Example 1 through a series reactor.

Figure 3 is a graph showing the changes in BTX yield over time in the one-step method in Comparative Example 1 and Comparative Example 2.

Figure 4 is a comparison of the BTX yield as a function of time with gallium-containing and non-gallium catalysts at 450°C based on ethane calculations in a series reactor.

Figure 5 shows the ethylene conversion rate, BTX yield, methane selectivity (a) and BTX component selectivity distribution (b) calculated based on ethylene over time in the first cycle run of the multi-cycle life experiment at 450°C and 3 bar. The change.

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

Detailed ways

The endpoints of ranges and any values disclosed herein are not limited to the precise range or value, but these ranges or values are to be understood to include values approaching such ranges or values. For numerical ranges, the endpoint values of each range, the endpoint values of each range and individual point values, and the individual point values can be combined with each other to obtain one or more new numerical ranges. These values The scope shall be deemed to be specifically disclosed herein.

Figure 1 is a simplified process flow diagram showing the method of the present invention. As shown in Figure 1, light hydrocarbons (light alkanes, boiling point does not exceed -10°C) such as ethane will first undergo dehydrogenation to generate corresponding olefins such as ethylene, and then the olefins will undergo oligomerization/aromatization to generate aromatic hydrocarbons and Other liquid fuels.

Regarding this two-step process, the inventor of the present invention has mentioned in previous patent documents that the conversion of lower alkanes into aromatic compounds is divided into two steps, in which the goal of step 1) is to dehydrogenate the alkane to obtain the progress of step 2). The goal of step 2) is to achieve high production volume and system efficiency of aromatic hydrocarbons and other liquid fuels through catalytic conversion of olefin-containing materials.

The present invention will continue and expand this 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 long-term catalyst stability, Optimize process conditions and catalyst performance. In addition, the product composition is not limited to aromatic compounds and may also include hydrocarbons in the gasoline range. The method can be designed to operate in different modes, such as targeting aromatics maximization, gasoline maximization, and overall catalyst performance and system efficiency optimization. The choice of operating mode can depend on market, demand and economics.

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

In the present invention, step 2) is preferably operated under lower temperature conditions, preferably 300-600°C, further preferably 350-600°C, and more preferably 400-550°C. The inventor of the present invention found that higher temperatures are conducive to the formation of BTX, but at the same time, the generation of CH ₄ and heavy aromatic hydrocarbon products with carbon atoms greater than 10 (A10+) also increases significantly, resulting in light aromatic hydrocarbons (A6-A9) The carbon utilization reduction is defined and calculated as the sum of the product amounts of lower carbon hydrocarbons (C2-C5).

Specifically, when BTX is used as the target product, the temperature in step 2) is preferably 450-550°C.

When using other liquid fuels such as gasoline as the target product, the temperature in step 2) is preferably 350-450°C.

In the present invention, liquid fuel refers to C5 or higher hydrocarbon components, that is, gasoline components.

When maximizing overall catalyst performance and system efficiency is the goal, the temperature in step 2) is preferably 400-500°C.

Step 2) can be carried out at normal pressure or at a pressure of 1-5 bar.

The liquid hourly space velocity WHSV calculated as ethylene of the olefin-containing stream may be 0.5-10 g/g-cat/hr, preferably 0.75-3 g/g-cat/hr.

Step 2) can be carried out in commonly used fixed bed reactors, fluidized bed reactors, and moving bed reactors.

The catalyst in step 2) is a zeolite molecular sieve with or without active metal components and a binder used for catalyst shaping.

The active metal component may be a metal element from Group IIIA, Group VIII and Group VIB of the periodic table of elements, such as one or more of Ga, Fe, Ni, Ag and Mo, preferably Ga and Ni. one or more. Based on the total amount of the catalyst and calculated as elements, the content of the active metal component is 0.4-5% by weight, preferably 0.8-2.5% by weight, such as 0.4, 0.5, 0.8, 1, 1.2, 1.4, 2, 2.1, 2.2, 2.3, 2.4, 2.5% by weight.

In the present invention, the zeolite molecular sieve can be various molecular sieves with a zeolite structure. Preferably, the silicon-aluminum molar ratio of the zeolite molecular sieve is 5-300, and further preferably ZSM-5, ZSM-11, ZSM-12, One or more of ZSM-23, ZSM-35, Y-type zeolite, beta zeolite, ferrierite and mordenite.

Based on the total amount of the catalyst, the content of zeolite molecular sieve is preferably not less than 50% by weight, more preferably more than 70% by weight, such as 75-85% by weight.

The aromatization catalyst may be a powder or a shaped body. Preferably, the aromatization catalyst is a shaped body containing active metal, zeolite molecular sieve and binder. The binder can be various substances that can shape the zeolite molecular sieve into a shaped catalyst, for example, it can be one or more of silica, aluminum oxide, clay, aluminum phosphate, and zirconium oxide. Preferably, the weight ratio of zeolite molecular sieve and binder is 50-99:1-50, preferably 70-90:10-30, and further preferably 75-85:15-25.

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

The above-mentioned catalyst can be a commercial product or can be prepared by a known method. For example, first, an impregnation method, such as a saturated impregnation method, is used to load the active metal component onto the zeolite molecular sieve, and then a conventional extrusion method is used to extrude the zeolite molecular sieve and the binder into a shape, for example, into a strip, column, or Clover type.

According to a preferred embodiment of the present invention, the aromatization catalyst is synthesized by wet impregnation by mixing a zeolite molecular sieve powder with a nitrate solution of a binder such as boehmite, water and active metal components to form a paste. The paste is extruded, dried and calcined at 500-650°C for 8-12 hours. The calcined extruded bars are crushed and sieved, and particles of 20-40 mesh are taken for catalytic reaction.

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

Since the dehydrogenation reaction proceeds independently, the conditions for the dehydrogenation reaction can be set such that relatively high conversions (e.g., >50%) can be achieved at temperatures significantly lower than typical ethane cracker outlet temperatures (generally > 850°C), the dehydrogenation temperature is, for example, about 750°C. Due to the relatively low temperature, coking is negligible in the absence of steam in this zone, so steam does not need to be introduced. In order to promote the reduction of the partial pressure of alkanes to promote the conversion of light hydrocarbons, it is preferred to pass a diluent, such as nitrogen, into the reactor.

If no steam is used in the dehydrogenation reaction, the product stream can be used directly for the next step of oligomerization/aromatization reaction without cooling separation to remove moisture and reheating for reaction. This will save the cost of the entire process. Since the oligomerization/aromatization reactions also proceed independently, the product stream obtained by the dehydrogenation reaction (generally an alkane/olefin/hydrogen mixture) can be selectively converted into aromatic compounds by selecting appropriate catalysts and controlling reaction conditions. or gasoline products. Since this step only requires the oligomerization/aromatization reaction and does not require the dehydrogenation of low-carbon alkanes that requires high activation energy, the catalyst used does not need to contain noble metals with strong dehydrogenation functions such as Pt, which can reduce the hydrogenation of ethylene. Major side reactions include ethane and further hydrogenolysis into methane, thus the amount of methane generated will be significantly reduced.

In the present invention, the dehydrogenation reaction can be carried out in the presence of a dehydrogenation catalyst or without a dehydrogenation catalyst.

The dehydrogenation catalyst can be various catalysts with alkane dehydrogenation function. Preferably, the dehydrogenation catalyst is a supported catalyst. The supported catalyst contains a carrier and a metal group with dehydrogenation activity supported on the carrier. The carrier is an inorganic heat-resistant oxide without acidic centers.

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

According to an embodiment of the present invention, based on the total amount of the dehydrogenation catalyst, the content of the metal component with dehydrogenation activity is 0.01-2.0% by weight, preferably 0.02-0.2% by weight.

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 above-mentioned catalyst may be a commercial product or may be prepared by a known method.

As mentioned above, since the dehydrogenation reaction in the present invention is carried out independently, the dehydrogenation reaction can be carried out under conditions favorable to dehydrogenation. Preferably, when the dehydrogenation reaction is carried out in the presence of a dehydrogenation catalyst, the dehydrogenation reaction The reaction is carried out at a temperature below 900°C, preferably 650-850°C, more preferably 650-800°C. This temperature is lower than the general ethane cracking reaction temperature (usually greater than 850°C).

Preferably, the GHSV of the light hydrocarbon stream is 500-20000h ^-1 , preferably 800-5000h ^-1 .

According to another embodiment of the present 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 700-900°C. Reactant residence time is preferably 0.05-30 seconds. In the present invention, the residence time of the reactants refers to the time the reactants stay at the above-mentioned dehydrogenation reaction temperature of 700-900°C, that is, the time of the dehydrogenation reaction.

In order to promote the reduction of the partial pressure of alkanes to promote the conversion of light hydrocarbons, it is preferred to pass a diluent into the dehydrogenation reactor. The diluent may be, for example, nitrogen or other inert gases that do not adversely affect the reaction.

Since step 1) of the present invention is to dehydrogenate light hydrocarbons, and the target product is the corresponding olefin, the goal of the dehydrogenation reaction is to obtain as many olefins as possible, so the product of the dehydrogenation reaction is called "olefin-containing logistics”. The olefin-containing stream can be directly subjected to the oligomerization/aromatization reaction described in step 2) without separation, thereby saving the time required for cooling, separation and heating up again to reach the temperature required for the oligomerization/aromatization reaction. time, shorten the process, and greatly reduce the resulting costs. Therefore, 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 a reaction in which the olefin stream obtained by dehydrogenation in step 1) oligomerizes and aromatizes to form aromatic hydrocarbons. It should be noted that the raw materials for the oligomerization/aromatization reaction in step 2) can all come from step 1), or can be obtained by adjusting based on the product of step 1) as needed, preferably so that the ethylene content is not less than 20 Volume %, preferably 20-50 volume %.

According to the present invention, the hydrocarbons in the light hydrocarbon stream can be various substances that can undergo dehydrogenation reaction and oligomerization/aromatization reaction to generate aromatic hydrocarbons, for example, they can be various alkanes with a carbon number of not more than 5. Preferably, the ethane content in the light hydrocarbon stream is not less than 65% by volume, more preferably 75-100% by volume.

Preferably, the ethylene content in the stream mainly containing olefins is not less than 20% by volume, preferably 20-50% by volume.

Preferably, the dehydrogenation reaction in step 1) and the oligomerization/aromatization reaction in step 2) are performed in different zones of the same reactor, or in different reactors. In order to enable the reaction to proceed continuously and not cause the operation to be suspended due to deactivation of the catalyst, preferably, the dehydrogenation reaction in step 1) and the oligomerization/aromatization reaction in step 2) are respectively arranged in parallel. Two, so that after the catalyst is deactivated, the reaction can be cut into another reactor, and the original reactor can regenerate the deactivated catalyst. The reactor may be a fixed bed reactor or a fluidized bed reactor.

According to a preferred embodiment of the present invention, in order to obtain higher purity liquid fuels such as aromatic hydrocarbons and gasoline products and convert light hydrocarbons into target products as much as possible, preferably, as shown in Figure 1, the method further includes The liquid fuel-containing stream obtained in step 2) is subjected to gas-liquid separation to obtain a liquid-phase stream containing liquid fuel. The liquid-phase stream containing liquid fuel is discharged as a product or sent to a subsequent separation process.

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

Further preferably, the light hydrocarbon stream (C2-C4 components) is returned to step 1) as raw material for dehydrogenation reaction.

The present invention will be described in detail below through examples. In the following examples, the molecular sieves used in the oligomerization/aromatization catalyst are commercially available products from Zeolyst Company. Unless otherwise stated, the content of active metal components is the weight percentage relative to the molecular sieve.

In the following result calculations, there are two types based on the ethane feed amount (also called ethane feed amount) and the ethylene feed amount (also called ethylene feed amount). All definitions of "amount" are molar amounts of carbon.

The calculation formula for ethane conversion rate is: ethane conversion rate, % = 100% × (ethane input - ethane output after reaction) / ethane input

The calculation formula for ethylene conversion rate is: ethylene conversion rate, % = 100% × (ethylene input amount - ethylene output after reaction) / ethylene input amount

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

The yield of component Y is calculated as: conversion rate of ethane or ethylene × selectivity of component Y (based on ethane or ethylene)

In order to calculate the ethylene conversion capacity and BTX production capacity, in the obtained curves of ethylene conversion rate and BTX yield changing with time, several calculation lower limits (slightly "lower limits") are usually defined in terms of ethylene conversion rate or BTX yield. For example, the lower limit of the BTX yield calculated based on ethylene is defined as 30%. In the life experiment, the time from the start of the reaction to the "lower limit" can be obtained, that is, the single-pass catalyst life (simply "life") and the reaction time within this reaction period is calculated. Average conversion rate.

Ethylene conversion capacity is calculated as: lifetime to reach the defined lower limit × average conversion calculated based on ethane or ethylene during this reaction time × liquid time space velocity of the ethane or ethylene feed

The BTX production capacity is calculated as: lifetime to reach the defined lower limit × average BTX yield calculated based on ethane or ethylene during this reaction time × liquid hourly space rate of ethane or ethylene feed

The composition of the product is determined online or offline using gas chromatography.

Example 1

Using the process shown in Figure 1, a C ₂ H ₆ /N ₂ mixed gas with a volume ratio of 1.67:1 is fed into two series-connected dehydrogenation and oligomerization/aromatization reactors, and the dehydrogenation temperature is set at 750 ℃, the GHSV of ethane is 1000h ^-1 , the pressure is normal pressure, and the reaction is carried out without a catalyst. Under these conditions, the conversion of ethane is approximately 57% and the concentration of ethylene in the dehydrogenated product gas stream is approximately 22% by volume. The set temperatures of the oligomerization/aromatization fixed bed reactor are 400°C, 450°C, 500°C and 550°C respectively, WHSV=2.68gC ₂ H ₆ /g-cat·hr (equivalent to 1.6gC ₂ H ₄ / g-cat·hr), the pressure is normal pressure (1 bar), and the catalyst is a Ga/ZSM-5/Al ₂ O ₃ catalyst prepared by a saturated impregnation method (Ga content is 2.0 wt%, ZSM-5 silicon aluminum The molar ratio is 30, and the weight ratio of molecular sieve to 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.

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

As can be seen from Figure 2(b), the yield of BTX is also greatly affected by temperature. The higher the temperature, the higher the initial yield of BTX, but it also decreases faster due to coking. The running times when the BTX yield dropped to the lower limit of 15% at these four temperatures were: 2705 minutes at 400°C, 3772 minutes at 450°C, 2747 minutes at 500°C and 1494 minutes at 550°C.

Figure 2(c) shows the _CH4 selectivity under four temperature conditions. The _CH4 selectivity measured at 400°C is about 8%. Due to the tandem mode operation, most of this 8% _CH4 comes from step 1, with the added portion being newly formed in step 2 oligomerization/aromatization. Corresponding to the generation of BTX, the initial selectivity of methane is high under high temperature conditions, but it also decreases rapidly as the catalyst deactivates.

In BTX products, the benzene fraction is greatly affected by temperature, as shown in (d) in Figure 2. 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. If the light aromatic hydrocarbon BTX is used as the target product, benzene has a higher value. However, when gasoline is the target product, the benzene content is strictly limited, and a low value is preferred.

Calculated based on the curve of BTX yield changing with time (Figure 2-b), the BTX production capacity with 15% BTX yield as the lower limit is 22.5 at 400°C, 40.3 at 450°C, 34.0 at 500°C and 18.4 at 550°C. g-BTX/g-cat, see Table 1.

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

Comparative example 1

The conversion of ethane was carried out according to the method of Example 1, except that a Pt/ZSM-5 catalyst was used for one-step ethane aromatization. Specifically, a mixture of ethane and nitrogen is used as the feed. The dehydrogenation/oligomerization/aromatization temperature is set at 630°C, WHSV=1.34gC ₂ H ₆ /g-cat.hr, the pressure is normal pressure, and the catalyst is Pt/ZSM-5 with a Pt loading of 0.05% by weight. /Al ₂ O ₃ (the molar ratio of silicon to aluminum of ZSM-5 is 30, and the weight ratio of molecular sieve to binder Al ₂ O ₃ is 70/30). The results are shown in Table 1 and (a) in Figure 3. It can be seen from (a) in Figure 3 that although the WHSV is only half of that of Example 1, the BTX yield decreases rapidly with the reaction time. A single cycle operation can only last about 400 minutes, and the BTX yield has dropped to The lower limit is 15%.

Comparative example 2

The conversion of ethane was carried out according to the method of Comparative Example 1, except that a Pt-Zn-Sn/ZSM-5 catalyst was used for one-step ethane aromatization. Specifically, a mixture of ethane and nitrogen is used as the feed. The dehydrogenation/oligomerization/aromatization temperature is set at 630°C, WHSV=1.34gC ₂ H ₆ /g-cat.hr, the pressure is normal pressure, the catalyst has a Pt loading of 0.05% by weight, and a Zn loading of 0.017 wt%, Sn loading is 0.03 wt% Pt-Zn-Sn/ZSM-5/Al ₂ O ₃ (the silicon-aluminum molar ratio of ZSM-5 is 30, and the weight ratio of molecular sieve to binder Al ₂ O ₃ is 70 /30). The results are shown in Table 1 and (b) in Figure 3. As can be seen from (b) in Figure 3, the BTX yield decreases more slowly than (a) in Figure 3, but the reaction can only run for 400 minutes.

In addition, comparing the single-cycle BTX production capacity of the two types of operations in Example 1 and Comparative Examples 1 and 2, it was found that the value of the 1-step method (Comparative Examples 1 and 2) was not higher than 5g-BTX/g-cat; One-step method, in the temperature range of 400-550℃, the single-cycle BTX production capacity is 18-40g-BTX/g-cat, which is several times that of the one-step method, see Table 1.

Comparative example 3

Carry out the conversion of ethane according to the method of Example 1, except that the ethylene aromatization catalyst used is ZSM-5 without metal addition (ZSM-5 is the same as that of Example 1, the weight of molecular sieve and binder Al ₂ O ₃ The ratio is 82.5/17.5), and the reaction is only carried out at a temperature of 450°C. The results of BTX yield changing with time are shown in Figure 4. Calculated based on this curve, the BTX production capacity with 15% BTX yield as the lower limit is 16.5g-BTX/g-cat.

Table 1

Example 2

Different from the method of Example 1, a mixture gas stream prepared according to the product composition of step 1) in Example 1 is used, and its composition and volume ratio are C ₂ H ₄ /H ₂ /N ₂ =1:1:1. The catalyst is the same as the oligomerization/aromatization catalyst in Example 1, that is, the Ga/ZSM-5/Al ₂ O ₃ catalyst prepared by the saturation impregnation method (the content of Ga is 2.0% by weight, the silicon and aluminum moles of ZSM-5 The ratio is 30, and the weight ratio of molecular sieve to binder Al ₂ O ₃ is 82.5/17.5). The conditions for aromatization are 450°C and 550°C, normal pressure, WHSV=0.75gC ₂ H ₄ /g-cat·hr, and the reaction time is 6 hours. The results of the 6-hour average BTX yield and product selectivity calculated based on ethylene are shown in Table 2.

Comparative example 4

The experiment was carried out according to the method of Example 2, except that the catalyst used for aromatization of ethylene was ZSM-5 without metal addition (ZSM-5 is the same as Example 1). The average results of the reaction at 450°C for 6 hours are shown in Table 2.

Example 3

The experiment was carried out according to the method of Example 2, except that the catalyst used for aromatization of ethylene was ZSM-5 containing 1% nickel of the total weight of the molecular sieve (ZSM-5 is the same as Example 1). The average results of the reaction at 450°C for 6 hours are shown in Table 2.

Example 4

The experiment was carried out according to the method of Example 2, except that the catalyst used for aromatization of ethylene was ZSM-5 containing 1% iron by total weight of molecular sieves (ZSM-5 is the same as Example 1). The average results of the reaction at 450°C for 6 hours are shown in Table 2.

Example 5

The experiment was carried out according to the method of Example 2, except that the catalyst used for aromatization of ethylene was ZSM-5 containing 1% silver of the total weight of molecular sieves (ZSM-5 is the same as Example 1). The average results of the reaction at 450°C for 6 hours are shown in Table 2.

Table 2

It can be seen from the data of Example 2 in Table 2 above that at 550°C, the BTX yield can be as high as 63.5%; the selectivities of CH ₄ and heavy aromatics (A10+) are 11.7% and 9.0% respectively. The collective selectivity of C2-C5 and A6-A9 can be calculated by adding up. According to the carbon utilization rate defined previously, that is, (C2-C5) + (A6-A9), then under this reaction condition, the carbon utilization rate About 80%.

At 450°C, the BTX yield is 55.1%, 8.4% lower than 550°C, but the selectivity for CH ₄ and A10+ is also lower, and the carbon utilization rate is about 90%, which is about 10% higher than 550°C.

Since the components in C2-C5 are mainly light hydrocarbons such as ethane and propane, these components can be returned to the dehydrogenation reactor through the circulation path shown in Figure 1 to produce more liquid products, resulting in high carbon utilization. This is a factor that also needs to be considered in the optimization of process conditions.

In the liquid product, the benzene fraction at 550°C is significantly higher than the value at 450°C, so if a high benzene content is targeted, the reactor needs to be operated at high temperatures.

It can be seen from the comparison results in Table 2 and Figure 4 that under the conditions of 450°C, in the catalytic reaction without gallium addition (Comparative Example 3, Comparative Example 4), it can be seen from Figure 2 that the BTX yields are respectively lower than those in the Examples 1 and Example 2, and it can be seen from the table that the selectivity of propane, butane, etc. among the products is significantly higher than the former (i.e., Example 2), thus revealing that the main effect of gallium is to make propane, butane, etc. Aromatize alkane to increase BTX yield.

Example 6

The experiment was carried out according to the method of Example 2. The difference is that the conditions for aromatization of ethylene are 450°C, 3 bar, WHSV=1.5gC ₂ H ₄ /g-cat·hr, and the volume ratio is 0.67:1:1:1 Gas mixture of C ₂ H ₆ /C ₂ H ₄ /H ₂ /N ₂ . The reaction was subjected to a 3-month multi-cycle life test. Each cycle run includes approximately 100 hours of reaction and approximately 24 hours of catalyst regeneration. Catalyst regeneration uses an air/nitrogen gas mixture with a volume ratio of 50:50.

For each cycle, the 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 testing were completed, a total of 1990g-C ₂ H ₄ /g-cat was converted (at the lower limit of 80% ethylene conversion rate), and 858g of BTX product was generated. On average per cycle, the ethylene conversion capacity is 117g-C ₂ H ₄ /g-cat and the BTX production capacity is 50g-BTX/g-cat.

Figure 5 reveals the changes in ethylene conversion, BTX yield, and methane selectivity over time (a) and the selectivity of each component in BTX over time (b) in the first catalytic cycle. In the completed life experiment of nearly 100 hours, the calculated ethylene conversion capacity and BTX production capacities are 127g-C ₂ H ₄ /g-cat (lower limit 1) and 140g-C ₂ H ₄ /g-cat (lower limit 1), 62g-BTX/g-cat (lower limit 2) and 66g-BTX /g-cat (lower limit 2).

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

The above examples fully illustrate that since the two reaction steps can be operated under respective optimized conditions, the two-step method provided by the present invention has obvious technical advantages. Specifically, step 1) can be an industrially mature ethane cracking process, step 2) can be operated in a lower temperature range, such as 350-550°C, and appropriate process conditions can be selected according to the target product: (1) ) Taking BTX, especially benzene, as the target product, step 2) can be carried out in the temperature range of 500-550°C; (2) taking gasoline products as the target product, the optimal temperature range of step 2) is between 350-400°C time; (3) aiming at the highest BTX yield, catalyst life and carbon utilization, then step 2) is performed at about 450°C to obtain the optimal performance of the overall catalyst and efficiency.

In addition, low-temperature operation also has other benefits. From an engineering point of view, lower-temperature operation and relatively long catalyst life will allow the use of simple and easy fixed-bed reactors, and the requirements for reactor materials are relatively low. All 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 concept of the present invention, many simple modifications can be made to the technical solution of the present invention, including the combination of various technical features in any other suitable manner. These simple modifications and combinations should also be regarded as the disclosed content of the present invention. All belong to the protection scope of the present invention.

Claims (19)

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;

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 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.