CN104818042A - Moving bed methanol-to-hydrocarbon method - Google Patents

Abstract

The invention relates to a moving bed methanol-to-hydrocarbon method. The method comprises the steps of hydrocarbon synthesis, separation and post-treatment, at least two serially connected reactors are adopted in the hydrocarbon synthesis step, a reaction raw material and a catalyst sequentially flow through the reactors in a countercurrent direction, low carbon olefin-containing circulation gas generated in the separation step returns to different feeding positions in the hydrogen synthesis step as quenching gas or raw material supplement gas in the hydrocarbon synthesis step, and a raw material methanol is used to wash and absorb C1-C4 light components generated in the separation step, returns, is fed and is converted in order to convert methanol into mixed aromatic hydrocarbon-containing stable light hydrocarbon with high added values. The method allows step complete utilization of the activity of the catalyst to be carried out, so the method improves the fine control of the reaction process, realizes effective material utilization and heat integration between processing processes, improves the product yield, reduces energy consumption and reduces environmental pollution.

Description

Moving-bed Methanol hydrocarbon method

Technical field

The present invention relates to a kind of methyl alcohol process for producing hydrocarbons adopting moving-bed.

Background technology

BTX aromatic hydrocarbons (Benzene, Toluene, Xylene) is the important basic raw material of petrochemical complex, and wherein p-Xylol (PX) demand is maximum.Along with the rapid expansion of domestic PX downstream PTA, production of polyester ability, on market, PX supplies wretched insufficiency, and to 2013, China's p-Xylol external dependence degree was up to 55.3%, and insufficiency of supply-demand strengthens further.Traditional technology production PX projects construction difficulty is large, production technology threshold is high, investment large, limits more by raw material naphtha resource.The increase that is nervous and consumers demand of current China's oil resource causes the shortage of resources such as raw material petroleum naphtha, solar oil of producing aromatic hydrocarbons, must seek new way and substitute traditional petroleum path production aromatic hydrocarbon product.What form sharp contrast therewith is domestic rich coal resources, is mainly that the methyl alcohol production capacity of raw material production is seriously superfluous with coal.In conjunction with the fundamental realities of the country of China's " oil starvation, weak breath, rich coal ", utilize abundant coal resources synthesizing methanol, research and development methanol oxidation transforms prepares aromatic hydrocarbons (MTA) technique, just high density PX can be obtained at production link, improve the added value of Downstream Products of Methanol, thus effectively reduce aromatic hydrocarbon product to the dependency of oil.

The aromatization of methanol technology of research and development both at home and abroad just progressively enters the industrialization stage at present, and portion of techniques realizes industrialization.MOBILE fixed bed Methanol aromatic hydrocarbons (gasoline processed) technology in 20th century the seventies achieve industrialization, and obtain industrial application at home; Shanxi coalification institute of Chinese Academy of Sciences bed technology obtained industrial application at home in 2010; Tsing-Hua University's fluidized-bed aromatization of methanol technology achieved ton industrial demonstration unit and runs in 2013.At present, fixed bed production technology range of application is comparatively wide, but is limited to the switching between reaction regeneration, and production capacity is restricted; Although fluidized-bed relies on the process of its successive reaction regeneration, production capacity has very large development space, the fluidization operation for this special material of methyl alcohol still needs to explore technique and operating method further.All there is certain shortcoming in current fixed bed and fluidized bed process mode, governs the extensive development in Methanol aromatic hydrocarbons field to some extent, specific as follows:

1) shortcoming of fixed bed operation mode:

(1) reaction regeneration frequently switches, and decaying catalyst needs to be interrupted regeneration, and reactor was significantly compressed for the time of reacting, production capacity critical constraints; (2) reaction regeneration frequently switches not only complex operation, and there is mishandle hidden danger, is unfavorable for long-term operation; (3) need for some time just can reach smooth running state by after regeneration incision reaction, material loss is larger; (4) general facilities consumption is large, and particularly reaction regeneration handoff procedure needs to consume a large amount of nitrogen; (5) easily there is the situation such as channel, bias current in production process in fixed bed, easily coking in reactor, and catalyzer duct easily blocks, and affects quality product and production safety; (6) fixed bed reaction heat removes difficulty, and catalyst change cost is high.

2) shortcoming of fluidized bed process mode:

(1) fluidized-bed layer inner catalyst back-mixing degree is heavier, and local reaction excessively easily causes coking; (2) in fluidized-bed layer, turbulence is violent, serious wear, and expensive catalyzer cracky and then generation are run and damaged, and cause loss economically; (3) in fluidized-bed layer, residence time destribution is comparatively wide, easily causes product slates wider, and the yield of target product reduces; (4) temperature and pressure surge all can affect the efficiency of gas solid separation system, and then affect subsequent fractionation system; (5) for the reactive system that coking yield is low, the reaction-regeneration system thermal equilibrium of fluidized-bed is difficult to maintain.

Summary of the invention

In order to overcome the above-mentioned defect under prior art, the object of the present invention is to provide a kind of moving-bed Methanol hydrocarbon method, the method can realize the serialization of aromatization of methanol reaction and catalyst regeneration process, the refinement controlling extent of reaction process can be improved, realizing between complete processing material effectively utilizes with heat integrated, the advantage such as have that catalyst activity is stable, pressure drop is low, plug flow reaction, back-mixing are few.

Technical scheme of the present invention is:

A kind of moving-bed Methanol hydrocarbon method, comprise hydrocarbon synthesis step, at least two reactors of mutually connecting are adopted in described hydrocarbon synthesis step, anti-applications catalyst regenerates according to entering revivifier by most top reactor successively after each reactor to the order of least significant end reactor, then most top reactor is returned, methanol feedstock is introduced into least significant end reactor after heating up, its reaction product enters its last reactor as reaction raw materials, the rest may be inferred, until the reaction product of the second reactor enters most top reactor as reaction raw materials.Described reactor is radially moving bed reactor, and can be " π " type reactor or " Z " type reactor, can be to cardioid reactor or centrifugal type reactor.

Described moving-bed Methanol hydrocarbon method also comprises separating step and post-processing step, described separating step successively adopts gas-oil-water three-phase separating device and single-stage or the product of multistage fractionation plant to described hydrocarbon synthesis step to carry out separation and Extraction, the reaction product of most top reactor sends into described gas-oil-water three-phase separating device afterwards according to processing requirement cooling (such as to 40 ~ 60 DEG C), be separated the most of gas phase obtained and be pressurized to 0.25 ~ 1.9MPaG as circulation gas through recycle gas compressor compression, described circulation gas sub-thread enters described least significant end reactor as reaction raw materials after heating up, or enter most top reactor and least significant end reactor respectively as reaction raw materials after being divided into two strands to heat up separately, by regulating the reaction depth of each stock circulation gas Flow-rate adjustment respective reaction device, consider that the circulation gas entering most top reactor improves except yield except participating in reaction, also to play cold shock effect, therefore the circulation gas flow entering most top reactor is preferably made to be greater than the circulation gas flow entering least significant end reactor, the mode that circulation gas enters respective reaction device is after heating up and the reaction feed of respective reaction device is converged and entered respective reaction device again, be separated the small portion gas phase obtained and enter post-processing step as pending material, be separated the aqueous portion that obtains and send into oil-contained waste water treatment device, also can direct reuse to upstream coal gasification apparatus, thus effectively save general facilities, be separated the oil phase part obtained and be distributed into described fractionation plant, the liquid phase C6 that fractionation obtains ~ C10 component or C8 ~ C10 component are drawn as product, correspondingly, liquid phase C3 ~ C5 component or C3 ~ C7 component return hydrocarbon synthesis step as reaction raw materials, converge respectively with different reactor charging, all the other light constituents enter post-processing step as pending material.Together post-processing step is introduced after the small portion gas phase that separation obtains and all the other light constituents that fractionation obtains can converge.The type of cooling of the reaction product of most top reactor can be the combination of dry type air cooling, wet type air cooling, water-cooled or aforesaid way.

The reaction product of most top reactor is preferably divided into multiply before the described gas-oil-water three-phase separating device of feeding, and the methanol feedstock and per share circulation gas that are about to enter hydrocarbon synthesis step are heated up respectively by heat exchange as exothermic medium, after reaction product cooling, multiply is converged and is cooled.The flow carrying out the reaction product of heat exchange with methanol feedstock is preferably less than the flow of the reaction product of carrying out heat exchange with circulation gas, and, when carrying out heat exchange with multiply circulation gas, the flow carrying out the reaction product of heat exchange with the circulation gas entering least significant end reactor is preferably less than the flow carrying out the reaction product of heat exchange with the circulation gas entering most top reactor, to meet the heating needs to more circulation gas.

Described post-processing step methyl alcohol is treated treated substance and is carried out reverse normal temperature washing, equipment adopts absorption tower, methanol feedstock self-absorption tower top enters absorption tower, from top to down is to by enter bottom absorption tower and the pending material risen sprays, and liquid at the bottom of absorption tower sends into described least significant end reactor as reaction raw materials after heating up; Absorb tower top non-condensable gas to be discharged by tower top, enter bleed-off system and use as fuel gas, or enter methanol-water cleaning device in order to reclaim methyl alcohol.Methanol absorption tower top non-condensable gas contains part methyl alcohol, is equipped with methanol recovery device at its downstream direction.Reclaim methyl alcohol can reuse to upstream coal gasification apparatus, also can be used as combustion gas.

The fractionation plant of described single-stage is preferably depentanizer or separation column, when adopting depentanizer, makes C6 ~ C10 aromatic hydrocarbons mixing prod enter product storage tank by discharging at the bottom of tower; Top gaseous phase C1 ~ C5 cools through the combination type of cooling of dry type air cooling, wet type air cooling, water-cooled or aforesaid way, temperature is down between the boiling point of C4 and C5 under logistics current pressure, enter return tank of top of the tower, C1 ~ C4 gaseous component is discharged by the tank deck of return tank of top of the tower, the light constituent obtained as fractionation enters bottom absorption tower, C5 liquid phase is through the supercharging of trim the top of column pump, and part backflow returns tower top; Another part returns most top reactor after heating up as reaction raw materials; When adopting separation column, C8 ~ C10 aromatic hydrocarbons mixing prod is made to enter product storage tank by discharging at the bottom of tower; Top gaseous phase C1 ~ C7 cools through the combination type of cooling of dry type air cooling, wet type air cooling, water-cooled or aforesaid way, temperature is down between the boiling point of C5 and C6 under logistics current pressure, enter return tank of top of the tower, C1 ~ C5 gaseous component is discharged by the tank deck of return tank of top of the tower, the light constituent obtained as fractionation enters bottom absorption tower, C6 ~ C7 liquid phase is through the supercharging of trim the top of column pump, and part backflow returns tower top; Another part C6 ~ C7 liquid phase as product extraction, or returns most top reactor after heating up as reaction raw materials, or part as product extraction, another part returns most top reactor after heating up as reaction raw materials.Enter most top reactor after mixing with the reaction product (i.e. intermediates) of next reactor of most top reactor after C5 or the C6 ~ C7 returning most top reactor heats up to participate in reacting.

Described multistage fractionation plant preferably includes depentanizer and debutanizing tower, gas-oil-water three-phase separating device is separated the oil phase part obtained and is distributed into described depentanizer, C6 ~ C10 aromatic hydrocarbons mixing prod is made to enter product storage tank by discharging at the bottom of depentanizer tower, depentanizer top gaseous phase C1 ~ C5 enters depentanizer top return tank and is separated into gas phase and liquid phase after condensation, gas phase is discharged by the tank deck of depentanize return tank of top of the tower, liquid phase is through the supercharging of depentanize tower top reflux pump, part backflow returns depentanizer tower top, and another part enters debutanizing tower, the condensation of depentanizer top gaseous phase and be separated method be following any one: (1) depentanizer top gaseous phase temperature is reduced between the boiling point of C2 and C3 under logistics current pressure, isolates C1 ~ C2 gas phase and C3 ~ C5 liquid phase, (2) depentanizer top gaseous phase temperature is reduced between the boiling point of C4 and C5 under logistics current pressure, isolates C1 ~ C4 gas phase and C5 (usually also containing a small amount of C1 ~ C4) liquid phase, through debutanizing tower fractionation, C5 liquid-phase product is discharged by the bottom of debutanizing tower tower, most top reactor is returned through heating up as reaction raw materials, debutanizing tower top gaseous phase C1 ~ C4 enters debutanizing tower top return tank and is separated into gas phase and liquid phase after condensation, gas phase is discharged by the tank deck of debutylize return tank of top of the tower, liquid phase is through the supercharging of debutylize tower top reflux pump, part backflow returns debutanizing tower tower top, another part returns least significant end reactor after heating up as reaction raw materials, the gaseous component of discharging from the tank deck of depentanizer top return tank and debutylize return tank of top of the tower enters post-processing step as pending material, namely enter bottom absorption tower, the condensation of debutanizing tower top gaseous phase and be separated method be following any one: (1) debutanizing tower top gaseous phase temperature is reduced between the boiling point of C2 and C3 under logistics current pressure, isolates C1 ~ C2 gas phase and C3 ~ C4 liquid phase, (2) debutanizing tower top gaseous phase temperature is reduced between the boiling point of C3 and C4 under logistics current pressure, isolates C1 ~ C3 gas phase and C4 (usually containing a small amount of C3) liquid phase.Depentanizer top gaseous phase and debutanizing tower top gaseous phase condensing mode are the combination type of cooling cooling of dry type air cooling, wet type air cooling, water-cooled or aforesaid way.Enter most top reactor after mixing with the reaction product (i.e. intermediates) of next reactor of most top reactor after the C5 returning most top reactor heats up to participate in reacting.

Should cool and the method be separated according to the above-mentioned suitable top gaseous phase of actually operating pressure selection in practice, such as, when depentanizer/debutanizing tower working pressure is higher, as 1.5MPaG, corresponding method (1) should be selected, when depentanizer/debutanizing tower working pressure is lower, as 0.4MPaG, corresponding method (2) should be selected, with avoid due to different components boiling point too close to and affect separating effect, ensure good separating effect.

Described separating step can also adopt dehydrogenation reactor, in this case, C3 ~ C4 liquid phase (corresponding aforesaid method (1)) of debutylize trim the top of column or the C4 liquid phase (corresponding aforesaid method (2)) containing part C3 are through the supercharging of debutylize tower top reflux pump, part backflow returns debutanizing tower tower top, another part enters dehydrogenation reactor dehydrogenation after heating up, and C3, C4 unsaturated hydrocarbons obtained after dehydrogenation returns most top reactor after heating up as reaction raw materials.C3 ~ C4 liquid phase before dehydrogenation and C3, C4 unsaturated hydrocarbons after dehydrogenation all realize heating up by heat exchange, C3 ~ C4 liquid phase before dehydrogenation is preferably warming up to 350 ~ 540 DEG C, C3, C4 unsaturated hydrocarbons after dehydrogenation is preferably warming up to 150 ~ 250 DEG C, enter most top reactor after mixing with the reaction product (i.e. intermediates) of next reactor of most top reactor to participate in reacting, intensification thermal source can be the reaction product of each reactor or outer supplying heat source.

When adopting the fractionation plant of single-stage, preferably adopt following processing parameter: fresh methanol charging 1 pumping outside battery limit (BL) in hydrocarbon synthesis step, boosts to 0.2 ~ 1.8MPaG, and temperature is 25 ~ 40 DEG C.Liquid hourly space velocity in each reactor is 1 ~ 5h ^-1; The regeneration temperature of revivifier is 500 ~ 650 DEG C, and regeneration pressure is 0.2 ~ 1.9MPaG; The pressure of most top reactor is 0.12 ~ 1.73MPaG, and temperature is 370 ~ 550 DEG C; The pressure of least significant end reactor is 0.15 ~ 1.75MPaG, and temperature is 320 ~ 520 DEG C; Often in adjacent two reactors the top pressure of last reactor not higher than the top pressure of a rear reactor, and all not higher than the top pressure of least significant end reactor, the minimal pressure of last reactor not higher than the minimal pressure of a rear reactor, and all not higher than the minimal pressure of least significant end reactor; Sub-thread circulation gas is warming up to 320 ~ 480 DEG C; The circulation gas entering most top reactor is warming up to 250 ~ 480 DEG C, and the circulation gas entering least significant end reactor is warming up to 320 ~ 480 DEG C; Methanol feedstock is warming up to 250 ~ 480 DEG C, and first the mixing with methanol feedstock of liquid at the bottom of absorption tower together heats up with methanol feedstock again.Methyl alcohol spray flow and bottom gas phase rising throughput ratio are 5 ~ 20, and service temperature is normal temperature, and pressure is 0.3 ~ 1.4MPaG; The tower top pressure of depentanizer is 0.3 ~ 1.75MPaG, and tower reactor pressure is 0.35 ~ 1.8MPaG; The tower top pressure of separation column is 0.06 ~ 1.6MPaG, such as 0.06,0.8,1.4MPaG, can determine according to practical situation, tower reactor pressure is 0.07 ~ 1.8MPaG, such as 0.07,0.85,1.5MPaG, can determine according to practical situation; C5 or the C6 ~ C7 returning most top reactor realizes heating up by heat exchange, and be warming up to 150 ~ 250 DEG C, thermal source is the reaction product of each reactor or outer supplying heat source.

When adopting multistage fractionation plant, preferably adopt following processing parameter: fresh methanol charging 1 pumping outside battery limit (BL) in hydrocarbon synthesis step, boosts to 0.2 ~ 1.8MPaG, and temperature is 25 ~ 40 DEG C.Liquid hourly space velocity in each reactor is 1 ~ 5h ^-1, the regeneration temperature of revivifier is 500 ~ 650 DEG C, and regeneration pressure is 0.2 ~ 1.9MPaG; The pressure of most top reactor is 0.2 ~ 1.73MPaG, and temperature is 370 ~ 550 DEG C; The pressure of least significant end reactor is 0.25 ~ 1.75MPaG, and temperature is 320 ~ 520 DEG C; Often in adjacent two reactors the top pressure of last reactor not higher than the top pressure of a rear reactor, and all not higher than the top pressure of least significant end reactor, the minimal pressure of last reactor not higher than the minimal pressure of a rear reactor, and all not higher than the minimal pressure of least significant end reactor; Sub-thread circulation gas is warming up to 320 ~ 480 DEG C; The circulation gas entering most top reactor is warming up to 250 ~ 480 DEG C, and the circulation gas entering least significant end reactor is warming up to 320 ~ 480 DEG C; Methanol feedstock is warming up to 250 ~ 480 DEG C, and first the mixing with methanol feedstock of liquid at the bottom of absorption tower together heats up with methanol feedstock again.Methyl alcohol spray flow and bottom gas phase rising throughput ratio are 5 ~ 20, and service temperature is normal temperature, and pressure is 0.3 ~ 1.4MPaG; The tower top pressure of depentanizer is 0.3 ~ 1.75MPaG, and tower reactor pressure is 0.35 ~ 1.8MPaG; The tower top pressure of debutanizing tower is 0.4 ~ 1.6MPaG, and tower reactor pressure is 0.45 ~ 1.65MPaG; The C5 liquid phase returning most top reactor realizes heating up by heat exchange, be warming up to 150 ~ 250 DEG C, thermal source is the reaction product of each reactor or outer supplying heat source, and C3 ~ C4 liquid phase first the mixing with methanol feedstock returning least significant end reactor together heats up with methanol feedstock again.

For aforementioned moving-bed Methanol hydrocarbon method described in any one, multiply is divided into after methanol feedstock can also being heated up, except wherein one enters except least significant end reactor, other each stocks do not enter other reactors, and make the methanol feedstock accounting entering least significant end reactor be greater than the methanol feedstock entering other each reactors.The alkane such as methyl alcohol directly adds reactor, can be reaction and provides CH3-group, the C5 that the LPG returned with circulation gas, later separation part return carry out the reaction such as alkane aromatization, methanol alkylation, promote the carrying out of aromatization, are conducive to improving aromatics yield.

Above-mentioned heat exchange all can be realized by the interchanger of more than 1 or 2 serial or parallel connection.Its thermal source can be the reaction product of certain reactor, also can be outer supplying heat source.

Beneficial effect of the present invention is:

1, adopt moving-bed to carry out aromatization of methanol, overcome the shortcomings such as fixed bed production capacity is low, pressure drop is large, catalyst life is short, the easy coking and blocking of bed; Overcome again the shortcomings such as fluidized-bed back-mixing degree is large, catalyzer is easy to wear, race damage.Utilize moving-bed successive reaction to regenerate, ensure that increasing substantially of production capacity; Be heated for methyl alcohol easily decompose, the feature of easily coking completely in the short period of time, utilize movable bed catalyst plug flow to move, the radial contact reacts of raw material, effectively ensure that methanol conversion and reaction degree of uniformity; Utilize the feature that radially moving bed pressure drop is low, effectively saved energy waste, catalyzer plug flow in bed moves down, two-phase transportation, and flow velocity is low, avoids the wearing and tearing of catalyzer, effectively controls the distribution of reaction product, improves the selectivity of target product.While guarantee methyl alcohol high conversion, improve product yield.

2, adopting the form of multiple reactors in series, flows in inverse order in reaction raw materials flow direction and catalyst motion direction.Fresh methanol is heated and is easily decomposed, and first enter the least significant end reactor of arranged in series, react with the catalyst exposure through its penultimate reactor pre-passivating, lower catalyst activity efficiently avoid methyl alcohol short period of time decomposes.React in least significant end reactor and generate the intermediates such as low-carbon (LC) hydro carbons based on methyl alcohol, methanol feedstock provides CH3-group in addition, also promotes aromatization of methanol, the alkylating generation of hydro carbons to a certain extent.Under lower catalytic type activity, reaction temperature and, reaction temperature rising is less, reaction be easy to control.Reaction intermediate enters most top reactor, with the high activated catalyst contact reacts from revivifier.Because catalyst activity is higher, effectively can be promoted based on the more difficult reaction carried out of aromatization of low carbon hydrocarbon, hydrocarbon restructuring etc., when part methanol feedstock enters most top reactor in addition, provided CH3-group, also promoted the degree that alkylation transforms.Therefore speed of reaction is very fast, and reaction efficiency is high, is conducive to the generation of aromatic hydrocarbons target product.

Temperature rise is there is larger relative to single reactor operation, operation controls the problems such as difficulty is larger, this technique adopts multiple reactors in series, reaction raw materials and catalyzer are against order direction flow pattern, utilize the feature that raw material reaction speed is different and different to catalyst activity sexual demand, efficiently avoid methyl alcohol and cross thermolysis, both carried out utilizing completely to the high low activity of catalyzer, carry out utilizing step by step to it according to reaction depth again, the refinement achieving reaction process controls, effectively control reaction temperature rising, the complexity that improve differential responses process mates adaptability with catalyst activity height, promote while being conducive to product purity and yield.

3, the application makes full use of the feature that reaction product potential temperature is high, latent heat is large of each reactor, with its heat, one or many preheating is carried out to the circulation gas that reaction feed and product separation go out, efficiently utilize own heat and achieve the up to standard of reaction raw materials temperature, thus save outer heat supplied.

Adopt and reaction product is divided into multiply and the mode of difference preheated feed and circulation gas, can by regulating the throughput ratio of multiply logistics, flexible feeding temperature, makes feeding temperature and reaction temperature rising match.Thus make whole reaction have very strong regulating power and anti-fluctuation ability.This Energy Efficient that can realize is recycled, and the hot integration mode that can realize again flexible has saved energy effectively.

4, the gas phase portion that reaction product obtains after vapour, oil, water three phase separation pressurizes as circulation gas through compressor, is divided into two stocks not return 2 reactor feeds, and has played different effects respectively:

(1) circulation gas 1: return least significant end reactor, namely mix with fresh methanol charging 1, due in least significant end reactor to generate lower carbon number hydrocarbons intermediate product, aromatization of methanol, the CH3-group that C1 ~ C4 component in circulation gas can provide with methyl alcohol participates in reacting jointly, improve yield, promote that alkylation transforms.

(2) circulation gas 2: return most top reactor, namely mix with the reaction product of second reactor.Due in most top reactor with alkane aromatization, be reassembled as master, catalyst activity is high, and reaction is violent, and heat release is large, and therefore circulation gas plays cooling/cold shock effect to most top reactor, prevents from reacting too fast coking.Methyl alcohol in the alkane that C1 ~ C4 component in circulation gas and later separation part return and raw material 2 facilitates methanol alkylation and reacts, and is conducive to the generation promoting aromatic hydrocarbons, improves PX selectivity.

By regulating two strands of circulation gas flows, the reaction depth of adjustable most top reactor and least significant end reactor, enables the effectively relay of two reactor reactions, coupling, improves aromatics yield.

5, after reaction product three phase separation, liquid phase component enters depentanizer, obtains C6 ~ C10 target product at the bottom of tower.Tower top C1 ~ C5 component is through cooling, gas-liquid separation, and C5 component returns most top reactor feed, namely mixes with the reaction product of second reactor.Because most top catalyst reactor activity is high, be swift in response, there is again CH3-group, the lighter hydrocarbons that C5 component and circulation gas return can generate aromatic hydrocarbons by aromatization rapidly in most top reactor, thus efficiently utilize the value of C5 byproduct, decrease whole device byproduct quantity, improve aromatics yield.

6, depentanize tower top light constituent C1 ~ C4 enters in an absorption tower, draws one methanol feeding 2 couples of C1 ~ C4 and carries out spray-absorption, effectively C3, C4 component in depentanizer top gas is absorbed and is dissolved in wherein from fresh methanol charging.This strand of material mixes with the charging entering least significant end reactor, namely mixes with fresh methanol charging 1, thus efficiently utilizes C3, C4 component, utilizes the solvability of raw material self, adds quantity and the diversity of raw material, decreases raw material consumption.This normal temperature methanol wash column mode can realize absorbing efficiently at normal temperatures, has both eliminated the reboiler of conventional fractionation tower height energy consumption, and the raw material that make use of again technique self, as absorbing medium, all has huge advantage with creative from energy-conservation with conservation aspect.

In addition, the non-condensable gas in depentanizer top gas washes out by normal temperature methanol wash column mode, the hydrogen-containing gas particularly in system, effectively prevent non-condensable gas gathering in systems in which.

7, the setting of debutanizing tower significantly reduces the load of depentanize tower top, and depentanize tower top condensing temperature only need higher than the boiling temperature of C2.C5 and C1 ~ C4 is separated by debutanizing tower, C5 component returns most top reactor feed, because most top catalyst reactor activity is high, be swift in response, there is again CH3-group, the lighter hydrocarbons that C5 component and circulation gas return can generate aromatic hydrocarbons by aromatization rapidly in most top reactor, thus efficiently utilize the value of C5 byproduct, decrease whole device byproduct quantity, improve aromatics yield.Tower top C1, C2 component and C3, C4 Component seperation, C1, C2 component and depentanizer C1, C2 component enter normal temperature methanol wash column absorption tower after converging.Debutanizing tower effectively reduces C3, C4 content entering absorption tower, decreases methanol usage, achieves the saving of raw material and the energy.

8, debutanizing tower is separated the C3 ~ C4 component obtained and is converted into C3, C4 unsaturated hydrocarbons by the setting of dehydrogenation reactor effectively, converge with C5 at the bottom of debutanizing tower tower, and then mix with intermediates, enter reactor 1, the C=C himself contained participates in reaction, participate in cyclisation, restructuring directly, add product yield.In addition, the setting of dehydrogenation reactor improves the hydrogen purity of methanol wash column absorption tower tower top non-condensable gas effectively, and high hydrogen purity is conducive to the separation of subsequent gases.

9, three phase separation tank is separated one charging as normal temperature methanol wash column absorption tower of gas phase extraction obtained, and in time by the hydrogen extraction in system, the hydrogen effectively reduced in system is assembled.

10, this process products only comprises stable light hydrocarbon containing BTX aromatics and small part non-condensable gas, does not produce liquefied gas, the whole recycle of byproduct, farthest achieves effective utilization of material.By the setting to separation column operating parameters, can the light aromatic hydrocarbons of by-product part C6, C7, according to the market conditions flexible adjusting device product category of aromatic hydrocarbon product kind.

Accompanying drawing explanation

Fig. 1 is the general flow chart of first embodiment of the present invention;

Fig. 2 is the general flow chart of second embodiment of the present invention;

Fig. 3 is the general flow chart of the 3rd embodiment of the present invention;

Fig. 4 is the general flow chart of the 4th embodiment of the present invention.

Embodiment

The invention provides a kind of moving-bed Methanol hydrocarbon method, describe the method utilization aborning in detail below by way of several specific embodiment.

Embodiment one (see Fig. 1): containing the 1st, the 2 two reactor, separating step adopts depentanizer fractionation.

Fresh methanol charging pumping outside battery limit (BL), boosts to 0.5MPaG, temperature 25 DEG C.First fresh methanol charging 1 enters the 2nd reactor (being equivalent to least significant end reactor) after heating up with reaction product heat exchange, carry out radially moving bed contact reacts with from the 1st reactor (being equivalent to most top reactor) through the catalyzer of pre-passivating, liquid hourly space velocity is 1.0h ^-1, generate intermediates (i.e. the reaction product of the 2nd reactor), pressure 0.45MPaG, temperature 520 DEG C.Enter the 1st reactor after 2nd reactor product intermediates leave, carry out radially moving bed contact reacts with the high activated catalyst from revivifier, liquid hourly space velocity is 1.0h ^-1, formation reaction product, pressure 0.42MPaG, temperature 550 DEG C.Reaction product is divided into 2 strands after being drawn by the 1st reactor---and reaction product 1, reaction product 2, throughput ratio is 0.52.Reaction product 1 and methanol feeding 1 carry out heat exchange in the 1st heat exchange unit, and methanol feeding 1 is heated to 480 DEG C.Reaction product 2 carries out heat exchange with the circulation gas from recycle gas compressor in the 2nd heat exchange unit, after reaction product 1 after heat exchange is converged with reaction product 2, mode through dry type air cooling and water-cooled is cooled to 40 DEG C, enters the three phase separation that three phase separation tank carries out gas, oil, water.

Catalyzer is promoted to regenerator overhead, falls in revivifier and regenerate after leaving the 2nd reactor, regeneration temperature 650 DEG C, regeneration pressure 0.6MPaG.This revivifier is conventional regeneration device.High activated catalyst after revivifier regeneration is promoted to the 1st reactor head, carries out moving bed radial contact reacts, then enters the 2nd reactor, carry out moving bed radial contact reacts with the 2nd reactor feed with the reaction product from the 2nd reactor.

Gaseous component after the three phase separation that three phase separation tank carries out gas, oil, water is divided into two strands: gas phase 1, gas phase 2, and throughput ratio is 18.0.Gas phase 1 enters recycle gas compressor, is pressurized to 0.52MPaG.Circulation gas and the reaction product 2 of leaving recycle gas compressor carry out heat exchange in the 2nd heat exchange unit, and circulation gas is heated to 320 ~ 480 DEG C, such as 480 DEG C, mix, jointly as the reaction feed of the 2nd reactor with the methanol feeding 1 after heating up.

Oil phase component after the three phase separation that three phase separation tank carries out gas, oil, water enters depentanizer.Depentanizer operating parameters is as follows: tower top pressure: 0.5MPaG; Tower reactor pressure: 0.55MPaG.Through depentanizer fractionation, in liquid-phase reaction product, below C5 component (i.e. C1 ~ C5) is discharged by tower top, and C6 ~ C10 aromatic hydrocarbons mixing prod enters product storage tank by discharging at the bottom of tower.Depentanizer top gaseous phase is through dry type air cooling, the cooling of the water-cooled combination type of cooling, and temperature is down to 50 DEG C, enters depentanize return tank of top of the tower.C1 ~ C4 gaseous component is discharged by tank deck, and C5 liquid phase is through the supercharging of depentanize tower top reflux pump, and part backflow returns depentanizer tower top; Another part C5 liquid-phase product is warming up to 250 DEG C through the 5th heat exchange unit, returns the 1st reactor feed, namely mixes with the 2nd reactor product intermediates and participates in reacting as reaction feed, and the 5th heat exchange unit thermal source is outer for 1.2MPaG steam.

The gas phase 2 that C1 ~ C4 gaseous component of being discharged by depentanize return tank of top of the tower top is separated with three phase separation tank enters bottom absorption tower after converging.Absorption tower adopts normal temperature methanol wash column operating method, and one fresh methanol charging 2 (25 DEG C) enters absorption tower by tower top, sprays from top to bottom.Top, absorption tower methyl alcohol spray flow and bottom C1 ~ C4 gas phase rising throughput ratio are 20.0, tower top service temperature 25 DEG C, working pressure 0.3MPaG.Absorb tower top non-condensable gas (C1, C2 component) to be discharged by tower top; Liquid phase at the bottom of tower mixes with fresh methanol charging 1 before reactor, participates in reaction as reaction feed.The non-condensable gas absorbing tower top discharge enters follow-up methanol-water cleaning device, in order to reclaim methyl alcohol.The methanol waste water obtained returns coal gasification unit after treatment.

Embodiment two (see Fig. 2): containing the 1st, the 2 two reactor, separating step adopts depentanizer, the fractionation of debutanizing tower two-stage.

Fresh methanol charging pumping outside battery limit (BL), boosts to 1.76MPaG, temperature 30 DEG C.First fresh methanol charging 1 enters the 2nd reactor (being equivalent to least significant end reactor) after heating up with reaction product heat exchange, carry out radially moving bed contact reacts with from the 1st reactor (being equivalent to most top reactor) through the catalyzer of pre-passivating, liquid hourly space velocity is 4.8h ^-1, also can be 5.0h ^-1, generate intermediates (i.e. the reaction product of the 2nd reactor), pressure is 1.74MPaG or 1.75MPaG, and temperature is 350 DEG C or 320 DEG C.Enter the 1st reactor after 2nd reactor product leaves, carry out radially moving bed contact reacts with the high activated catalyst from revivifier, air speed is 2.0h ^-1, formation reaction product, pressure 1.72MPaG or 1.73MPaG, temperature 380 DEG C, also can be 370 DEG C.Reaction product is divided into 2 strands after being drawn by the 1st reactor---and reaction product 1, reaction product 2, throughput ratio is: 1.3.Reaction product 1 and methanol feeding 1 carry out heat exchange in the 1st heat exchange unit, and methanol feeding 1 is heated to 250 or 270 DEG C.Reaction product 2 is divided into reaction product 3, reaction product 4, and throughput ratio is 0.92, carries out heat exchange respectively with the circulation gas from recycle gas compressor in the 2nd heat exchange unit, the 3rd heat exchange unit.Reaction product 1 after heat exchange, reaction product 3, reaction product 4 are converged, and through being cooled to 50 DEG C, enter the three phase separation that three phase separation tank carries out gas, oil, water.

Catalyzer is promoted to regenerator overhead, falls in revivifier and regenerate after leaving the 2nd reactor, regeneration temperature 500 DEG C, regeneration pressure 1.85MPaG, also can be 1.9MPaG.This revivifier is conventional regeneration device.High activated catalyst after revivifier regeneration is promoted to the 1st reactor head, carries out moving bed radial contact reacts, then enters the 2nd reactor, carry out moving bed radial contact reacts with the 2nd reactor feed with the intermediates from the 2nd reactor.

Gaseous component after the three phase separation that three phase separation tank carries out gas, oil, water is divided into two strands: gas phase 1, gas phase 2, and throughput ratio is 15.0.Gas phase 1 enters recycle gas compressor, is pressurized to 1.84MPaG.The circulation gas leaving recycle gas compressor is divided into 2 strands---and circulation gas 1, circulation gas 2, throughput ratio is: 1.2.Circulation gas 1 and reaction product 3 carry out heat exchange in the 2nd heat exchange unit, and circulation gas is heated to 320 DEG C.Circulation gas 2 and reaction product 4 carry out heat exchange in the 3rd heat exchange unit, and circulation gas is heated to 350 DEG C.Circulation gas 1 mixes, jointly as the reaction feed of the 2nd reactor with the methanol feeding 1 after intensification; Circulation gas 2 mixes with the intermediates logistics of the 2nd reactor outlet, jointly as the reaction feed of the 1st reactor.

Oil phase component after the three phase separation that three phase separation tank carries out gas, oil, water enters depentanizer.Depentanizer operating parameters is as follows: tower top pressure: 1.75MPaG; Tower reactor pressure: 1.80MPaG.Through depentanizer fractionation, in liquid-phase reaction product, below C5 component is discharged by tower top, and C6 ~ C10 aromatic hydrocarbons mixing prod enters product storage tank by discharging at the bottom of tower.Depentanizer top gaseous phase is through the cooling of wet type air cooling, and temperature is down to 40 DEG C, enters depentanize return tank of top of the tower.C1 ~ C2 gaseous component is discharged by tank deck, and C3 ~ C5 liquid phase (usually also containing a small amount of C1, C2) is through the supercharging of depentanize tower top reflux pump, and part backflow returns depentanizer tower top; Another part enters debutanizing tower.

Debutanizing tower operating parameters is as follows: tower top pressure: 1.6MPaG; Tower reactor pressure: 1.65MPaG.Through debutanizing tower fractionation, top gaseous phase (C1 ~ C4 component) is discharged by tower top; C5 liquid-phase product at the bottom of tower is warming up to 160 DEG C through the 5th heat exchange unit, returns the 1st reactor feed, namely mixes with the 2nd reactor product intermediates and participates in reacting as reaction feed, and the 5th heat exchange unit thermal source is outer for 1.2MPaG steam.Debutanizing tower top gaseous phase is through the cooling of wet type air cooling, and temperature is down to 40 DEG C, enters debutylize return tank of top of the tower.C1 ~ C2 gaseous component is discharged by tank deck, and C3, C4 liquid phase is through the supercharging of debutylize tower top reflux pump, and part backflow returns debutanizing tower tower top; Another part liquid-phase product mixes with fresh methanol charging 1 before reactor, participates in reaction as reaction feed.

The gas phase 2 that depentanizer tower top C1 ~ C2 gaseous component is separated with debutanizing tower tower top C1 ~ C2 gaseous component and three phase separation tank enters bottom absorption tower after converging.Absorption tower adopts normal temperature methanol wash column operating method, and one fresh methanol charging 2 (30 DEG C) enters absorption tower by tower top, sprays from top to bottom.Top, absorption tower methyl alcohol spray flow and bottom C1 ~ C4 gas phase rising throughput ratio are 5 or 6.Tower top service temperature 25 DEG C, working pressure 1.3 or 1.4MPaG.Absorb tower top non-condensable gas (C1, C2 component) to be discharged by tower top; Liquid phase at the bottom of tower mixes with fresh methanol charging 1 before reactor, participates in reaction as reaction feed.The non-condensable gas absorbing tower top discharge enters follow-up methanol-water cleaning device, in order to reclaim methyl alcohol.The methanol waste water obtained returns coal gasification unit after treatment.

Embodiment three (see Fig. 3): containing the 1st, the 2 two reactor, separating step adopts depentanizer, the fractionation of debutanizing tower two-stage, and adopts dehydrogenation reactor to C3 ~ C4 dehydrogenation, and methanol feedstock divides 2 stocks not enter two reactors.

Fresh methanol charging pumping outside battery limit (BL), boosts to 0.3MPaG, temperature 30 DEG C.Fresh methanol charging 1 is divided into 2 strands after heating up with reaction product heat exchange: raw material 1 and raw material 2, respectively as the 1st reactor (being equivalent to most top reactor), the 2nd reactor (being equivalent to least significant end reactor) charging, throughput ratio is 1:9.Raw material 2 enters the 2nd reactor, and carry out radially moving bed contact reacts with from the 1st reactor through the catalyzer of pre-passivating, liquid hourly space velocity is 2.5h ^-1, generate intermediates (i.e. the reaction product of the 2nd reactor), pressure 0.25MPaG, temperature 420 DEG C.2nd reactor product intermediates leave afterwards and raw material 1 is mixed into the 1st reactor, and carry out radially moving bed contact reacts with the high activated catalyst from revivifier, liquid hourly space velocity is 2.5h ^-1, formation reaction product, pressure 0.2MPaG, temperature 460 DEG C.Reaction product is divided into 2 strands after being drawn by the 1st reactor---and reaction product 1, reaction product 2, throughput ratio is: 0.75.Reaction product 1 and methanol of reaction charging 1 carry out heat exchange in the 1st heat exchange unit, and methanol feeding 1 is heated to 365 DEG C.Reaction product 2 is divided into reaction product 3, reaction product 4, and throughput ratio is 0.76, carries out heat exchange respectively with the circulation gas from recycle gas compressor in the 2nd heat exchange unit, the 3rd heat exchange unit., the reaction product 1 after heat exchange, reaction product 3, reaction product 4 are converged, and through being cooled to 60 DEG C, enter the three phase separation that three phase separation tank carries out gas, oil, water.

Catalyzer is promoted to regenerator overhead, falls in revivifier and regenerate after leaving the 2nd reactor, regeneration temperature 570 DEG C, regeneration pressure 0.3MPaG, also can be 0.2MPaG.This revivifier is conventional regeneration device.High activated catalyst after revivifier regeneration is promoted to the 1st reactor head, carries out moving bed radial contact reacts, then enters the 2nd reactor, carry out moving bed radial contact reacts with the 2nd reactor feed with the intermediates from the 2nd reactor.

Gaseous component after the three phase separation that three phase separation tank carries out vapour, oil, water is divided into two strands: gas phase 1, gas phase 2, and throughput ratio is 12.0.Gas phase 1 enters recycle gas compressor, is pressurized to 0.3MPaG.The circulation gas leaving recycle gas compressor is divided into 2 strands---and circulation gas 1, circulation gas 2, throughput ratio is 0.9.Circulation gas 1 and reaction product 3 carry out heat exchange in the 2nd heat exchange unit, and circulation gas is heated to 365 DEG C.Circulation gas 2 and reaction product 4 carry out heat exchange in the 3rd heat exchange unit, and circulation gas is heated to 420 DEG C.Circulation gas 1 mixes, jointly as the reaction feed of the 2nd reactor with the raw material 2 after intensification; Intermediates and the raw material 1 of circulation gas 2 and the 2nd reactor outlet mix, jointly as the reaction feed of the 1st reactor.

Oil phase component after the three phase separation that three phase separation tank carries out vapour, oil, water enters depentanizer.Depentanizer operating parameters is as follows: tower top pressure: 0.3MPaG; Tower reactor pressure: 0.35MPaG.Through depentanizer fractionation, in liquid-phase reaction product, below C5 component is discharged by tower top, and C6 ~ C10 aromatic hydrocarbons mixing prod enters product storage tank by discharging at the bottom of tower.Depentanizer top gaseous phase is through the cooling of wet type air cooling, and temperature is down to 45 DEG C, enters depentanize return tank of top of the tower.C1 ~ C4 gaseous component is discharged by tank deck, and C5 liquid phase (usually also containing a small amount of C1 ~ C4) is through the supercharging of depentanize tower top reflux pump, and part backflow returns depentanizer tower top; Another part liquid-phase product enters debutanizing tower.

Debutanizing tower operating parameters is as follows: tower top pressure: 0.4 or 0.5MPaG; Tower reactor pressure: 0.45 or 0.55MPaG.Through debutanizing tower fractionation, top gaseous phase (C1 ~ C4 component) is discharged by tower top; C5 liquid-phase product at the bottom of tower is warming up to 180 DEG C through the 5th heat exchange unit, returns the 1st reactor feed, namely mixes with the 2nd reactor product intermediates and raw material 1 and participates in reacting as reaction feed, and the 5th heat exchange unit thermal source is outer for 1.2MPaG steam.Debutanizing tower top gaseous phase is through the cooling of wet type air cooling, and temperature is down to 40 DEG C, enters debutylize return tank of top of the tower.C1 ~ C3 gaseous component is discharged by tank deck, and C4 (containing a small amount of C3) liquid phase is through the supercharging of debutylize tower top reflux pump, and part backflow returns debutanizing tower tower top; Another part liquid-phase product, through the 4th heat exchange unit and reaction product 1 heat exchange, is heated to 430 DEG C, enters dehydrogenation reactor.Liquid-phase mixing at the bottom of dehydrogenation reactor outlet streams and debutanizing tower, and then mix with intermediates and raw material 1, participate in reaction as the 1st reactor reaction charging.

The gas phase 2 that depentanizer tower top C1 ~ C4 gaseous component is separated with debutanizing tower tower top C1 ~ C3 gaseous component and three phase separation tank enters bottom absorption tower after converging.Absorption tower adopts normal temperature methanol wash column operating method, and one fresh methanol charging 2 (30 DEG C) enters absorption tower by tower top, sprays from top to bottom.Top, absorption tower methyl alcohol spray flow and bottom C1 ~ C4 gas phase rising throughput ratio are 7.1.Tower top service temperature 25 DEG C, working pressure 0.8MPaG.Absorb tower top non-condensable gas (C1, C2 component) to be discharged by tower top; Liquid phase at the bottom of tower mixes with fresh methanol charging 1 before reactor, participates in reaction as reaction feed.

Embodiment four (see Fig. 4): containing the 1st, the 2nd, the 3 three reactor, separating step adopts the fractionation of separation column single-stage, and methanol feedstock divides 3 stocks not enter three reactors.

Fresh methanol charging pumping outside battery limit (BL), boosts to 0.6MPaG, temperature 30 DEG C.Fresh methanol charging is divided into 3 strands after heating up with reaction product heat exchange: raw material 1, raw material 2, raw material 3, respectively as the 1st reactor (being equivalent to most top reactor), the 2nd reactor, the 3rd reactor (being equivalent to least significant end reactor) charging, throughput ratio is 1:2:8.Raw material 3 enters the 3rd reactor, and carry out radially moving bed contact reacts with from the 2nd reactor through the catalyzer of pre-passivating, liquid hourly space velocity is 1.6h ^-1, generate intermediates 1 (i.e. the reaction product of the 3rd reactor), pressure 0.55MPaG, temperature 470 DEG C.3rd reactor product intermediates 1 leave afterwards and raw material 2 is mixed into the 2nd reactor, and carry out radially moving bed contact reacts with the greater activity catalyzer from the 1st reactor, liquid hourly space velocity is 1.3h ^-1, generate intermediates 2 (i.e. the reaction product of the 2nd reactor), pressure 0.53MPaG, temperature 500 DEG C.2nd reactor product intermediates 2 are drawn afterwards and raw material 1 is mixed into the 1st reactor, and carry out radially moving bed contact reacts with the high activated catalyst from revivifier, liquid hourly space velocity is 1.1h ^-1, formation reaction product, pressure 0.51MPaG, temperature 510 DEG C.

Reaction product is divided into 2 strands after being drawn by the 1st reactor---and reaction product 1, reaction product 2, throughput ratio is: 1.8.Reaction product 1 and methanol of reaction charging 1 carry out heat exchange in the 1st heat exchange unit, and methanol feeding 1 is heated to 420 DEG C.Reaction product 2 is divided into reaction product 3, reaction product 4, reaction product 5, and throughput ratio is 6:4:2, carries out heat exchange respectively with from the circulation gas of recycle gas compressor and Fractionator Bottom product in the 2nd heat exchange unit, the 3rd heat exchange unit, the 5th heat exchange unit.Reaction product 1 after heat exchange, reaction product 3, reaction product 4, reaction product 5 are converged, and through being cooled to 40 DEG C, enter the three phase separation that three phase separation tank carries out gas, oil, water.

Catalyzer is promoted to regenerator overhead, falls in revivifier and regenerate after leaving the 3rd reactor, regeneration temperature 590 DEG C, regeneration pressure 0.6MPaG.This revivifier is conventional regeneration device.High activated catalyst after revivifier regeneration is promoted to the 1st reactor head, carries out moving bed radial contact reacts with the intermediates 2 from the 2nd reactor; Enter the 2nd reactor again, carry out moving bed radial contact reacts with the intermediates 1 from the 3rd reactor; Enter the 3rd reactor again, carry out moving bed radial contact reacts with the 3rd reactor feed.

Gaseous component after the three phase separation that three phase separation tank carries out vapour, oil, water is divided into two strands: gas phase 1, gas phase 2, and throughput ratio is 10.0.Gas phase 1 enters recycle gas compressor, is pressurized to 0.64MPaG.The circulation gas leaving recycle gas compressor is divided into 2 strands---and circulation gas 1, circulation gas 2, throughput ratio is 1.1.Circulation gas 1 and reaction product 3 carry out heat exchange in the 2nd heat exchange unit, and circulation gas is heated to 420 or 480 DEG C.Circulation gas 2 and reaction product 4 carry out heat exchange in the 3rd heat exchange unit, and circulation gas is heated to 470 or 480 DEG C.Circulation gas 1 mixes, jointly as the reaction feed of the 3rd reactor with the raw material 3 after intensification; Intermediates 2 and the raw material 1 of circulation gas 2 and the 2nd reactor outlet mix, jointly as the reaction feed of the 1st reactor.

Oil phase component after the three phase separation that three phase separation tank carries out gas, oil, water enters separation column, and separation column operating parameters is as follows: tower top pressure is 0.06 ~ 1.6MPaG, such as 0.8MPaG; Tower reactor pressure is 0.07 ~ 1.8MpaG, such as 0.85MPaG; Through separation column fractionation, below C7 component is discharged by tower top, and C8 ~ C10 aromatic hydrocarbons mixing prod enters product storage tank by discharging at the bottom of tower; Top gaseous phase C1 ~ C7 cools through the mode of wet type air cooling, and temperature is down to 130 DEG C, enters fractionation return tank of top of the tower.C1 ~ C5 gaseous component is discharged by tank deck, and the gas phase 2 be separated with three phase separation tank enters bottom absorption tower after converging.C6 ~ C7 liquid phase is through the supercharging of fractionation tower top reflux pump, and part backflow returns fractionator overhead; Another part is warming up to 190 DEG C after the 5th heat exchange unit and reaction product 5 heat exchange, returns the 1st reactor feed, namely mixes with the 2nd reactor product intermediates 2 and raw material 1 and participates in reacting as reaction feed.

The gas phase 2 that fractionator overhead C1 ~ C5 gaseous component is separated with three phase separation tank enters bottom absorption tower after converging.Absorption tower adopts normal temperature methanol wash column operating method, and one fresh methanol charging 2 (30 DEG C) enters absorption tower by tower top, sprays from top to bottom.Top, absorption tower methyl alcohol spray flow and bottom gas phase rising throughput ratio are 9.2.Tower top service temperature 25 DEG C, working pressure 0.8MPaG.Absorb tower top non-condensable gas (C1, C2 component) to be discharged by tower top; Liquid phase at the bottom of tower mixes with fresh methanol charging 1 before reactor, participates in reaction as reaction feed.

The present invention verifies according to embodiment 1,2,3,4, and the result obtained is as follows:

Table 1 reaction raw materials forms

Composition	Mol％
		Methyl alcohol	99.9
Water	0.1

Table 2 product forms

Composition	Embodiment 1 (Mol%)	Embodiment 2 (Mol%)	Embodiment 3 (Mol%)	Embodiment 4 (Mol%)
					C4	0.1	0.1	0.1	0.01
C5	0.56	0.55	0.61	0.06
					C6A	12.23	12.01	11.52	4.6
C7A	30.16	30.02	29.08	15.83
					C8A	43.04	42.76	43.4	53.69
C9A	11.46	12.19	12.53	19.74
					C10A	2.45	2.37	2.76	6.07

Table 3 non-condensable gas forms

Composition	Embodiment 1 (Mol%)	Embodiment 2 (Mol%)	Embodiment 3 (Mol%)	Embodiment 4 (Mol%)
					CO	2.23	2.19	0.74	2.21
Hydrogen	16.22	16.4	72.07	16.33
					H 2O	0	0	0	0
Methanol	2.39	2.01	0.80	2.28
					C1	58.13	57.96	19.38	58.01
C2	20.62	21.06	6.87	20.78
					C3	0.28	0.26	0.1	0.27
C4	0.13	0.12	0.04	0.12

The present invention changes traditional single reaction vessel the form of more than 2 reactors in series into, utilize the feature that raw material reaction speed is different and different to catalyst activity sexual demand, efficiently avoid methyl alcohol and cross thermolysis, violent reaction process be divided into and severally comparatively leniently react, the relay that is coupled successively is carried out.Both carried out utilizing completely to the high low activity of catalyzer, and carried out utilizing step by step again according to reaction depth to it, the refinement achieving reaction process controls, and effectively controls reaction temperature rising, promotes while being conducive to product purity and yield.Efficiently solve traditional single reaction vessel and operate the problems such as the temperature rise existed is comparatively large, operation control difficulty is larger.

Separating step is separated the C3 ~ C5 liquid-phase product obtained and turns back to reaction member participation reaction.Circulation gas returns and carries a large amount of CH3-groups, can react rapidly generation aromatic hydrocarbons, thus efficiently utilize the value of C5 byproduct in C5 Returning reactor, decreases whole device byproduct quantity.In C3 ~ C4 component Returning reactor, efficiently utilize C3, C4 component, effectively reduce raw material consumption, decrease the consumption of byproduct.

Claims (10)

1. a moving-bed Methanol hydrocarbon method, comprise hydrocarbon synthesis step, it is characterized in that in described hydrocarbon synthesis step, adopting at least two reactors of mutually connecting, anti-applications catalyst regenerates according to entering revivifier by most top reactor successively after each reactor to the order of least significant end reactor, then most top reactor is returned, methanol feedstock is introduced into least significant end reactor after heating up, its reaction product enters its last reactor as reaction raw materials, the rest may be inferred, until the reaction product of the second reactor enters most top reactor as reaction raw materials, described reactor is radially moving bed reactor.

2. moving-bed Methanol hydrocarbon method as claimed in claim 1, characterized by further comprising separating step and post-processing step, described separating step successively adopts gas-oil-water three-phase separating device and single-stage or the product of multistage fractionation plant to described hydrocarbon synthesis step to carry out separation and Extraction, described gas-oil-water three-phase separating device is sent into after the reaction product cooling of most top reactor, be separated the most of gas phase obtained and be used as circulation gas through recycle gas compressor compression, described circulation gas sub-thread enters described least significant end reactor as reaction raw materials after heating up, or enter most top reactor and least significant end reactor respectively as reaction raw materials after being divided into two strands to heat up separately, by regulating the reaction depth of each stock circulation gas Flow-rate adjustment respective reaction device, be separated the small portion gas phase obtained and enter post-processing step as pending material, be separated the aqueous portion obtained and send into oil-contained waste water treatment device, be separated the oil phase part obtained and be distributed into described fractionation plant, the liquid phase C6 that fractionation obtains ~ C10 component or C8 ~ C10 component are drawn as product, and correspondingly, liquid phase C3 ~ C5 component or C3 ~ C7 component return hydrocarbon synthesis step as reaction raw materials, and all the other light constituents enter post-processing step as pending material.

3. moving-bed Methanol hydrocarbon method as claimed in claim 2, it is characterized in that the reaction product of most top reactor is divided into multiply before the described gas-oil-water three-phase separating device of feeding, and the methanol feedstock and per share circulation gas that are about to enter hydrocarbon synthesis step are heated up respectively by heat exchange as exothermic medium, after reaction product cooling, multiply is converged.

4. moving-bed Methanol hydrocarbon method as claimed in claim 3, it is characterized in that described post-processing step methyl alcohol is treated treated substance and carried out reverse normal temperature washing, equipment adopts absorption tower, methanol feedstock self-absorption tower top enters absorption tower, from top to down is to by enter bottom absorption tower and the pending material risen sprays, and liquid at the bottom of absorption tower sends into described least significant end reactor as reaction raw materials after heating up; Absorb tower top non-condensable gas to be discharged by tower top, enter bleed-off system and use as fuel gas, or enter methanol-water cleaning device in order to reclaim methyl alcohol.

5. moving-bed Methanol hydrocarbon method as claimed in claim 4, is characterized in that the fractionation plant of described single-stage is depentanizer or separation column, when adopting depentanizer, makes C6 ~ C10 aromatic hydrocarbons mixing prod enter product storage tank by discharging at the bottom of tower; Top gaseous phase cools through the combination type of cooling of dry type air cooling, wet type air cooling, water-cooled or aforesaid way, temperature is down between the boiling point of C4 and C5 under logistics current pressure, enter return tank of top of the tower, C1 ~ C4 gaseous component is discharged by the tank deck of return tank of top of the tower, the light constituent obtained as fractionation enters bottom absorption tower, C5 liquid phase is through the supercharging of trim the top of column pump, and part backflow returns tower top; Another part returns most top reactor after heating up as reaction raw materials; When adopting separation column, C8 ~ C10 aromatic hydrocarbons mixing prod is made to enter product storage tank by discharging at the bottom of tower; Top gaseous phase cools through the combination type of cooling of dry type air cooling, wet type air cooling, water-cooled or aforesaid way, temperature is down between the boiling point of C5 and C6 under logistics current pressure, enter return tank of top of the tower, C1 ~ C5 gaseous component is discharged by the tank deck of return tank of top of the tower, the light constituent obtained as fractionation enters bottom absorption tower, C6 ~ C7 liquid phase is through the supercharging of trim the top of column pump, and part backflow returns tower top; Another part C6 ~ C7 liquid phase as product extraction, or returns most top reactor after heating up as reaction raw materials, or part as product extraction, another part returns most top reactor after heating up as reaction raw materials.

6. moving-bed Methanol hydrocarbon method as claimed in claim 4, it is characterized in that described multistage fractionation plant comprises depentanizer and debutanizing tower, gas-oil-water three-phase separating device is separated the oil phase part obtained and is distributed into described depentanizer, C6 ~ C10 aromatic hydrocarbons mixing prod is made to enter product storage tank by discharging at the bottom of depentanizer tower, depentanizer top gaseous phase enters depentanizer top return tank and is separated into gas phase and liquid phase after condensation, gas phase is discharged by the tank deck of depentanize return tank of top of the tower, liquid phase is through the supercharging of depentanize tower top reflux pump, part backflow returns depentanizer tower top, another part enters debutanizing tower, the condensation of depentanizer top gaseous phase and be separated method be following any one: (1) depentanizer top gaseous phase temperature is reduced between the boiling point of C2 and C3 under logistics current pressure, isolates C1 ~ C2 gas phase and C3 ~ C5 liquid phase, (2) depentanizer top gaseous phase temperature is reduced between the boiling point of C4 and C5 under logistics current pressure, isolates C1 ~ C4 gas phase and C5 liquid phase, through debutanizing tower fractionation, C5 liquid-phase product is discharged by the bottom of debutanizing tower tower, most top reactor is returned through heating up as reaction raw materials, debutanizing tower top gaseous phase enters debutanizing tower top return tank and is separated into gas phase and liquid phase after condensation, gas phase is discharged by the tank deck of debutylize return tank of top of the tower, liquid phase is through the supercharging of debutylize tower top reflux pump, part backflow returns debutanizing tower tower top, another part returns least significant end reactor after heating up as reaction raw materials, the gaseous component of discharging from the tank deck of depentanizer top return tank and debutylize return tank of top of the tower enters post-processing step, the condensation of debutanizing tower top gaseous phase and be separated method be following any one: (1) debutanizing tower top gaseous phase temperature is reduced between the boiling point of C2 and C3 under logistics current pressure, isolates C1 ~ C2 gas phase and C3 ~ C4 liquid phase, (2) debutanizing tower top gaseous phase temperature is reduced between the boiling point of C3 and C4 under logistics current pressure, isolates C1 ~ C3 gas phase and C4 liquid phase.

7. moving-bed Methanol hydrocarbon method as claimed in claim 6, it is characterized in that described separating step also adopts dehydrogenation reactor, in this case, the liquid phase of debutylize trim the top of column is through the supercharging of debutylize tower top reflux pump, part backflow returns debutanizing tower tower top, another part enters dehydrogenation reactor dehydrogenation after heating up, and the unsaturated hydrocarbons obtained after dehydrogenation returns most top reactor after heating up as reaction raw materials.

8. moving-bed Methanol hydrocarbon method as claimed in claim 5, is characterized in that the liquid hourly space velocity in each reactor is 1 ~ 5h ^-1, the regeneration temperature of revivifier is 500 ~ 650 DEG C, regeneration pressure is 0.2 ~ 1.9MPaG, the pressure of most top reactor is 0.20 ~ 1.73MPaG, temperature is 370 ~ 550 DEG C, the pressure of least significant end reactor is 0.25 ~ 1.75MPaG, temperature is 320 ~ 520 DEG C, often in adjacent two reactors the top pressure of last reactor not higher than the top pressure of a rear reactor, and all not higher than the top pressure of least significant end reactor, the minimal pressure of last reactor is not higher than the minimal pressure of a rear reactor, and all not higher than the minimal pressure of least significant end reactor, sub-thread circulation gas is warming up to 320 ~ 480 DEG C, the circulation gas entering most top reactor is warming up to 250 ~ 480 DEG C, the circulation gas entering least significant end reactor is warming up to 320 ~ 480 DEG C, methanol feedstock is warming up to 250 ~ 480 DEG C, the first mixing with methanol feedstock of liquid at the bottom of absorption tower together heats up with methanol feedstock again, methyl alcohol spray flow and bottom gas phase rising throughput ratio are 5-20, pressure is 0.3 ~ 1.4MPaG, the tower top pressure of depentanizer is 0.3 ~ 1.75MPaG, tower reactor pressure is 0.35 ~ 1.8MPaG, the tower top pressure of separation column is 0.06 ~ 1.6MPaG, tower reactor pressure is 0.07 ~ 1.8MpaG, C5 or the C6 ~ C7 returning most top reactor realizes heating up by heat exchange, be warming up to 150 ~ 250 DEG C, thermal source is the reaction product of each reactor or outer supplying heat source.

9. moving-bed Methanol hydrocarbon method as claimed in claims 6 or 7, is characterized in that the liquid hourly space velocity in each reactor is 1 ~ 5h ^-1, the regeneration temperature of revivifier is 500 ~ 650 DEG C, regeneration pressure is 0.2 ~ 1.9MPaG, the pressure of most top reactor is 0.20 ~ 1.73MPaG, temperature is 370 ~ 550 DEG C, the pressure of least significant end reactor is 0.25 ~ 1.75MPaG, temperature is 320 ~ 520 DEG C, often in adjacent two reactors the top pressure of last reactor not higher than the top pressure of a rear reactor, and all not higher than the top pressure of least significant end reactor, the minimal pressure of last reactor is not higher than the minimal pressure of a rear reactor, and all not higher than the minimal pressure of least significant end reactor, sub-thread circulation gas is warming up to 320 ~ 480 DEG C, the circulation gas entering most top reactor is warming up to 250 ~ 480 DEG C, the circulation gas entering least significant end reactor is warming up to 320 ~ 480 DEG C, methanol feedstock is warming up to 250 ~ 480 DEG C, the first mixing with methanol feedstock of liquid at the bottom of absorption tower together heats up with methanol feedstock again, methyl alcohol spray flow and bottom gas phase rising throughput ratio are 5 ~ 20, pressure is 0.3 ~ 1.4MPaG, the tower top pressure of depentanizer is 0.3 ~ 1.75MPaG, tower reactor pressure is 0.35 ~ 1.8MPaG, the tower top pressure of debutanizing tower is 0.4 ~ 1.6MPaG, tower reactor pressure is: 0.45 ~ 1.65MPaG, the C5 liquid phase returning most top reactor realizes heating up by heat exchange, be warming up to 150 ~ 250 DEG C, thermal source is the reaction product of each reactor or outer supplying heat source, C3 ~ C4 liquid phase first the mixing with methanol feedstock returning least significant end reactor together heats up with methanol feedstock again.

10. as the moving-bed Methanol hydrocarbon method in claim 1-9 as described in any one, it is characterized in that methanol feedstock is divided into multiply after heating up, except wherein one enters except least significant end reactor, other each stocks do not enter other reactors, and the methanol feedstock accounting entering least significant end reactor is greater than the methanol feedstock entering other each reactors.