CN104818043A - High efficiency heat integrated moving bed methanol aromatization method for coproducing liquefied gas - Google Patents

Abstract

The invention relates to a high efficiency heat integrated moving bed methanol aromatization method for coproducing liquefied gas. The method comprises the steps of hydrocarbon synthesis and separation, at least two serially connected reactors are adopted in the hydrocarbon synthesis step, reaction products of the reactors are used to respectively carry out heat exchange heating on reaction raw materials and low carbon olefin-containing circulation gas generated in the separation step in order to make the circulation gas return to different feeding positions in the hydrogen synthesis step as quenching gas or raw 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 make methanol covert into mixed aromatic hydrocarbons with high added values and to co-produce liquefied gas as a byproduct. The method allows heat exchange of the reaction products and step complete utilization of the activity of a 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

The High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas

Technical field

The present invention relates to a kind of methyl alcohol process for producing hydrocarbons adopting the co-production of liquefied gas of moving-bed, the heat of this technological reaction product can be circulated efficiently and be utilized.

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 High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas, 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 High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas, comprise hydrocarbon synthesis and separating 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 most top reactor after heating up, its reaction product enters a reactor thereafter as reaction raw materials, the rest may be inferred, until the reaction product of penultimate reactor enters least significant end reactor as reaction raw materials, described reactor is radially moving bed reactor, can be " π " type reactor or " Z " type reactor, can be to cardioid reactor or centrifugal type reactor.The reaction product of multiple reactor successively carries out heat exchange with methanol feedstock as exothermic medium, cascade raising temperature is carried out to methanol feedstock, the plurality of reactor at least comprises most top reactor and least significant end reactor, the reaction product of least significant end reactor is divided into multiply, wherein at least one for heating up to methanol feedstock, entered the temperature of charge of most top reactor according to the temperature of the reaction product of least significant end reactor by the flow control changing this burst of reaction product, and then control the temperature of reaction of most top reactor.Described separating step adopts gas-oil-water three-phase separating device to converge each stock and carries out three phase separation according to the most end reactor reaction product after processing requirement cooling (such as to 40 ~ 60 DEG C), and the type of cooling can be the combination of dry type air cooling, wet type air cooling, water-cooled or aforesaid way.Aqueous portion that three phase separation obtains sends 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 depentanizer separation and Extraction C6 ~ C10 aromatic hydrocarbon product; Be separated the small portion gas phase obtained and enter post-processing step, major part gas phase is pressurized to 0.25 ~ 1.9MPaG through recycle gas compressor compression and is used as circulation gas, wherein return the reactor of hydrocarbon synthesis step after most of circulation gas intensification as reaction raw materials, small portion circulation gas enters de-liquefied gas tower fractionation extraction C3 ~ C4 and draws as liquefied gas product; The C5 component that depentanizer and de-liquefied gas tower fractionate out returns hydrocarbon synthesis step as reaction raw materials, and all the other light constituents enter post-processing step as pending material.The reaction product of least significant end reactor preferably heats up to methanol feedstock prior to the reaction product of most top reactor, because the reaction product potential temperature of least significant end reactor is higher, latent heat is larger, utilize it to significantly improve the temperature of methanol feedstock, the temperature entering most top reactor for accurately controlling methanol feedstock reserves and regulates space more freely.When methanol feedstock only enters most top reactor, can control the temperature of reaction of most top reactor not higher than the temperature of reaction of least significant end reactor, in adjacent two reactors, the temperature of reaction of last reactor is not higher than the temperature of reaction of a rear reactor.

Most top reactor and least significant end reactor is entered respectively as reaction raw materials, by the reaction depth regulating these two strands of circulation gas flows independently to regulate most top reactor and least significant end reactor online after described most of circulation gas is divided into two strands to heat up separately.The flow returning the circulation gas of most top reactor is preferably greater than the flow of the circulation gas returning least significant end reactor, participates in reaction and improves product yield, and play cold shock/cooling effect to make more multi cycle gas.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.

Before each stock of the reaction product of least significant end reactor converges, also have two strands and heated up respectively to per share circulation gas by heat exchange as exothermic medium in multiply reaction product, after heat exchange cooling, each stock reaction product converges cooling again.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 most top reactor is preferably greater than the flow carrying out the reaction product of heat exchange with the circulation gas entering least significant end reactor, to meet the heating needs when the circulation gas entering most top reactor is more.

The method adopting de-liquefied gas tower fractionation to extract C3 ~ C4 be following any one: (1) makes C5 liquid phase discharge by the bottom of tower, returns least significant end reactor after heating up as reaction raw materials, top gaseous phase C1 ~ C4 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 C2 and C3 under logistics current pressure, enter return tank of top of the tower, wherein C1 ~ C2 gaseous component is discharged by the tank deck of return tank of top of the tower, enter post-processing step, C3 ~ C4 liquid phase is through the supercharging of trim the top of column pump, and part backflow returns tower top, another part is as the extraction of liquefied gas product, (2) make C5 liquid phase discharge by the bottom of tower, after heating up as reaction raw materials, return most top reactor, top gaseous phase C1 ~ C4 temperature is down between the boiling point of C3 and C4 under logistics current pressure, enter return tank of top of the tower, wherein C1 ~ C3 gaseous component is discharged by the tank deck of return tank of top of the tower, be cooled between the boiling point of C2 and C3 under logistics current pressure again through the supercharging of circulation residue gas compressor, enter de-liquefied gas tower top knockout drum, C1 ~ C2 gaseous component after separation is discharged by the tank deck taking off liquefied gas tower top knockout drum, enter post-processing step, C3 liquid-phase reflux after separation is to de-liquefied gas return tank of top of the tower, converge with C4 liquid phase component, C3 ~ C4 liquid phase is through the supercharging of trim the top of column pump, part backflow returns tower top, another part is as the extraction of liquefied gas product.In method (2), the cooling method of gas phase C1 ~ C4 and gas phase C1 ~ C3 can be the combination type of cooling of dry type air cooling, wet type air cooling, water-cooled or aforesaid way.Enter least significant end reactor again after mixing with the reaction product (i.e. intermediates) of penultimate reactor after returning the C5 intensification of least significant end reactor to participate in reacting.C1 ~ C2 gas phase together introduces post-processing step after can converging with being separated the small portion gas phase obtained.

C3 ~ C4 extracting method that should be suitable according to actually operating pressure selection in practice, such as, when de-liquefied gas tower working pressure is higher, as 1.5MPaG, suitable system of selection (1), when de-liquefied gas tower working pressure is lower, as 0.4MPaG, should system of selection (2), with avoid due to different components boiling point too close to and affect separating effect, ensure good separating effect.

The process of employing depentanizer separation and Extraction aromatic hydrocarbon product is: make 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, wherein C1 ~ C4 gaseous component is discharged by the tank deck of return tank of top of the tower, enter post-processing step, C5 liquid phase is through the supercharging of trim the top of column pump, and part backflow returns tower top; Another part returns least significant end reactor after heating up as reaction raw materials.Enter least significant end reactor again after mixing with the reaction product (i.e. intermediates) of penultimate reactor after returning the C5 intensification of least significant end reactor to participate in reacting.C1 ~ C4 gas phase together introduces post-processing step after can converging with being separated the small portion gas phase obtained.

Described post-processing step employing 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 to by enter bottom absorption tower and the pending material risen sprays, liquid at the bottom of absorption tower send into after heating up as reaction raw materials described in most top reactor; 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.

For the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of any one co-production of liquefied gas aforesaid, described small portion circulation gas first through residue gas compressor supercharging, or first can also be lowered the temperature again through pump supercharging through the cold heat exchange unit of ammonia before entering de-liquefied gas tower.Arrange supercharging device before entering de-liquefied gas tower in addition can ensure to meet the condition entering de-liquefied gas tower, do not rely on again the top hole pressure of recycle gas compressor simultaneously, thus facilitate the adjustment of many places logistics pressure.

When adopting the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of described co-production of liquefied gas to produce, 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 0.25 ~ 1.75MPaG, and temperature is 320 ~ 520 DEG C; The pressure of least significant end reactor is 0.2 ~ 1.73MPaG, and temperature is 370 ~ 550 DEG C; If more than two reactors, after often in adjacent two reactors, the top pressure of a reactor is not higher than the top pressure of last reactor, and all not higher than the top pressure of most top reactor, the minimal pressure of a rear reactor not higher than the minimal pressure of last reactor, and all not higher than the minimal pressure of most top reactor; 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 270 ~ 490 DEG C; Methanol feedstock is successively warming up to 98 ~ 250 DEG C and 250 ~ 480 DEG C; Liquid at the bottom of absorption tower first mixes again with methanol feedstock together cascade raising temperature with methanol feedstock; 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, tower reactor pressure is 0.35 ~ 1.8MPaG, the tower top pressure of de-liquefied gas tower is 0.3 ~ 1.6MPaG, tower reactor pressure is 0.35 ~ 1.65MPaG, the C5 returning least significant end reactor is warming up to 150 ~ 250 DEG C by heat exchange, the reaction product that thermal source is reactor described in hydrocarbon synthesis step or outer supplying heat source.

Gas-liquid separation device is provided with in the cold heat exchange unit of described ammonia, the cold heat exchange unit of described ammonia is separated into more than C3 liquid composition and C1 ~ C2 gaseous component by after described small portion circulation gas cooling, enter de-liquefied gas tower after the supercharging of more than C3 liquid composition, C1 ~ C2 gaseous component enters post-processing step; When described small portion circulation gas first through residue gas compressor supercharging, be pressurized to 0.5 ~ 1.8MPaG; When described small portion circulation gas is first through the cooling of ammonia cold heat exchange unit again through pump supercharging, is cooled to-13 ~ 30 DEG C, is pressurized to 0.5 ~ 1.8MPaG.

For the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of aforementioned co-production of liquefied gas described in any one, multiply is divided into after methanol feedstock can also being heated up, except wherein one enters except most top reactor, other each stocks do not enter other reactors, and make the methanol feedstock accounting entering most top 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 unit all can comprise 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, adopt the form of multiple reactors in series, reaction raw materials is identical through the order of reactor with catalyzer through the order of reactor.First fresh methanol enters the most top reactor of arranged in series, with the high activated catalyst contact reacts from revivifier.Consider that fresh methanol is heated easily to decompose, therefore keep the temperature of reaction that most top reactor is lower, make reaction process comparatively gentle, reaction temperature rising is less, effectively can avoid methyl alcohol short period of time decomposes, and reaction is easy to control.React in most top reactor except the reaction such as aromatization of methanol, hydrocarbon restructuring generates except aromatic hydrocarbons, due to lower temperature of reaction, also there is methyl alcohol and generate the side reactions such as intermediates such as low-carbon (LC) hydro carbons, the intermediates wider distribution thus generated.Because most top reactor catalyst activity is higher, while, transformation efficiency high at reaction efficiency is high, keep lower temperature of reaction, reaction can be made to be easy to control.

For 2 reactors in series, intermediate product temperature after heat exchange of the 1st reactor reduces, and continues to enter the 2nd reactor and the activity comparatively low catalyst contact reacts from the 1st reactor.Because intermediate product is containing part aromatic hydrocarbon product, therefore the alkane only needing lower catalyst activity the lower carbon number hydrocarbons in intermediates and later separation part to be returned continues aromatization, hydrocarbon recombinant conversion is aromatic product, in addition provide CH3-group when part methanol feedstock enters reactor, also promote the degree that alkylation transforms.Because catalyst activity is lower, reaction is comparatively gentle, and reaction temperature rising is less, effectively prevent coking, and reaction is easy to control.

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 equidirectional flow pattern, utilize the feature that raw material reaction speed is different and different to catalyst activity sexual demand, branch's utilization has been carried out to catalyst activity, 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, efficiently avoid methyl alcohol and cross thermolysis, make reactions steps all comparatively gentle, the refinement achieving reaction process controls, effectively control reaction temperature rising, promote while being conducive to product purity and yield.

3, the application utilizes the feature that reaction product potential temperature at different levels is high, latent heat is large, make full use of its heat and 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.In addition, in time the heat of intermediate product is removed, the reactor reaction temperature in its downstream is reduced, be conducive to keeping reaction process comparatively steady, the generation of slagging prevention.

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, and be divided at least 3 strands, wherein at least 2 stocks do not return 2 reactor feeds, and have played different effects respectively:

(1) circulation gas 1: return most top reactor, namely mix with fresh methanol charging 1, the CH3-group that the C1 ~ C4 component in circulation gas can provide with methyl alcohol participates in reacting jointly, improves yield, promotes that alkylation conversion and alkane aromatization transform.In addition, because most top reactor catalyst activity is high, reaction is violent, and heat release is large, and circulation gas 1 plays cooling/cold shock effect to most top reactor, prevents from reacting too fast coking.

(2) circulation gas 2: return least significant end reactor, the intermediates namely produced with penultimate reactor mix.Owing to turning to master with aromatization of low carbon hydrocarbon, alkyl in least significant end reactor, the methyl alcohol in the alkane that the C1 ~ C4 component in circulation gas and later separation part return and charging 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 product 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, C5 component returns least significant end reactor feed, namely mix with penultimate reactor product intermediates, participate in the hydrocarbon synthesis of least significant end reactor, effectively make use of the value of C5 byproduct, decrease whole device byproduct quantity, improve aromatics yield.For the situation also entering fresh methanol raw material to each reactor except most top reactor, because in least significant end reactor, methanol feeding brings CH3-group, C5 can generate aromatic hydrocarbons by rapid aromatization under catalyst action, thus effectively utilize the value of C5 byproduct further, reduce whole device byproduct quantity.

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 most top reactor, namely mixes with fresh methanol charging 1, thus efficiently utilizes C3, C4 component, utilize the solvability of raw material self, add quantity and the diversity of raw material, decrease 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, can effectively be separated the liquefied gas in component loops gas, C1 ~ C2 light constituent, C5 component by arranging de-liquefied gas tower, here the C5 component obtained is converged with the C5 component obtained by depentanizer and is returned reaction feed, efficiently utilizes reaction by-product; C1 ~ C2 lightweight gas mixes with depentanizer top gas and carries out normal temperature methanol wash column, reclaims C3 ~ C4 further as reaction feed; C3 ~ C4 liquefied gas component that de-liquefied gas tower produces exports as handicraft product liquefied gas, has enriched the product category of this technique.

8, by adopting residue gas compressor or the cold heat exchange unit of ammonia, even if when the pressure of described small portion circulation gas is lower than de-liquefied gas tower tower pressure (such as reaction pressure lower and de-liquefied gas tower height press operation), gas phase or the charging of liquid phase mode of de-liquefied gas tower also can be realized smoothly.Wherein, arranging the cold heat exchange unit of ammonia can when reaction pressure be lower, is effectively liquefied by circulation gas, then makes it reach to enter the condition of de-liquefied gas tower by supercharging, thus effectively from circulation gas, reclaims C3 ~ C4 component; And adopt the advantage of residue gas compressor gas-phase feed mode to be that its comprehensive energy consumption is lower.

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 has BTX aromatics and small part non-condensable gas, by-product liquefied gas, and most of intermediate product recycle to extinction utilizes, and farthest achieves effective utilization of material.

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 High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas, 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 and the fractionation of de-liquefied gas tower.

Fresh methanol charging 1 pumping outside battery limit (BL), boosts to 1.76MPaG, temperature 25 DEG C.First fresh methanol charging 1 enters the 1st reactor (being equivalent to most top reactor) after heating up with reaction product heat exchange, and carry out radially moving bed contact reacts with the highly active catalyzer from revivifier, liquid hourly space velocity is 5.0h ^-1, generate intermediates (i.e. the reaction product of the 1st reactor), pressure 1.75MPaG, temperature 350 DEG C, also can be 320 DEG C.Intermediates enter the 1st heat exchange unit after leaving the 1st reactor, and carry out preheating as thermal source to the 1st reactor feed, methanol feeding 1 is heated to 250 DEG C, and intermediates are cooled to 320 DEG C.Intermediates enter the 2nd reactor (being equivalent to least significant end reactor) after leaving the 1st heat exchange unit, carry out radially moving bed contact reacts with the catalyzer from the 1st reactor, and liquid hourly space velocity is 5.0h ^-1, formation reaction product, pressure 1.73MPaG, temperature 400 DEG C, also can be 370 DEG C.Reaction product is divided into 2 strands after being drawn by the 2nd reactor---and reaction product 1, reaction product 2, throughput ratio is 0.66.Reaction product 1 and methanol feeding 1 carry out heat exchange in the 2nd heat exchange unit, and methanol feeding 1 is heated to 98 DEG C.Reaction product 2 is divided into reaction product 3, reaction product 4, and throughput ratio is 0.96, carries out heat exchange respectively with the circulation gas from recycle gas compressor in the 3rd heat exchange unit, the 4th heat exchange unit.Reaction product 1 after heat exchange, reaction product 3, reaction product 4 are converged, and are cooled to 45 DEG C through wet type air cooling, 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 1.86 or 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 intermediate product from the 1st reactor with the 1st reactor feed.

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 1.84MPaG.The circulation gas leaving recycle gas compressor is divided into 3 strands---and circulation gas 1, circulation gas 2, circulation 3, throughput ratio is 0.9:1.0:1.0.Circulation gas 1 and reaction product 3 carry out heat exchange in the 3rd heat exchange unit, and circulation gas is heated to 250 DEG C or 270 DEG C.Circulation gas 2 and reaction product 4 carry out heat exchange in the 4th heat exchange unit, and circulation gas is heated to 270 or 320 DEG C.The fresh methanol charging 1 in circulation gas 1 and the 1st heat exchange unit heat-absorbing medium exit converges, jointly as the reaction feed of the 1st reactor; The intermediates that circulation gas 2 and the 1st heat exchange unit exothermic medium export converge, jointly as the reaction feed of the 2nd 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: 1.8MPaG.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 cools through the array mode of dry type air cooling, water-cooled, and temperature is down to 115 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 150 DEG C through the 5th heat exchange unit, mixes with the 1st reactor product intermediates, participates in reaction as the 2nd reactor reaction charging, and the 5th heat exchange unit thermal source is outer for 1.2MPaG steam.

Circulation gas 3 enters de-liquefied gas tower.De-liquefied gas tower operating parameters is as follows: tower top pressure: 1.6MPaG; Tower reactor pressure: 1.65MPaG.Through de-liquefied gas tower fractionation, in circulation gas 3, below C4 component is discharged by tower top, and C5 liquid-phase product is discharged by the bottom of tower, mixes with C5 liquid-phase product at the bottom of depentanizer, participates in reaction as the 2nd reactor reaction charging.De-liquefied gas column overhead gas phase cools through wet type air cooling mode, and temperature is down to 40 DEG C, enters de-liquefied gas return tank of top of the tower.C1 ~ C2 gaseous component is discharged by tank deck, and C3 ~ C4 liquid phase is through de-liquefied gas trim the top of column pump supercharging, and part backflow returns de-liquefied gas column overhead; Another part C3 ~ C4 liquid-phase product is as the extraction of liquefied gas product.

The gas phase 2 that the C1-C4 gaseous component of being discharged by depentanize return tank of top of the tower top is separated with de-liquefied gas column overhead 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 (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 5.The temperature on this absorption tower is normal temperature, working pressure 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 two (see Fig. 2): containing the 1st, the 2 two reactor, separating step adopts depentanizer and the fractionation of de-liquefied gas tower, circulation gas 3 is first through the cold heat exchange of ammonia and pump supercharging before entering de-liquefied gas tower, and methanol feedstock divides 2 stocks not enter two reactors.

Fresh methanol charging 1 pumping outside battery limit (BL), boosts to 0.3MPaG, temperature 25 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 9:1.Raw material 1 enters the 1st reactor, carries out radially moving bed contact reacts with the highly active catalyzer from revivifier, and liquid hourly space velocity is 1.0h ^-1, generate intermediates (i.e. the reaction product of the 1st reactor): pressure 0.25MPaG, temperature 520 DEG C.Intermediates enter the 1st heat exchange unit after leaving the 1st reactor, and carry out preheating as thermal source to the 1st reactor feed, methanol feeding 1 is heated to 480 DEG C, and intermediates are cooled to 490 DEG C.Be mixed into the 2nd reactor with raw material 2 after intermediates leave the 1st heat exchange unit, carry out radially moving bed contact reacts with the catalyzer from the 1st reactor, liquid hourly space velocity is 1.0h ^-1, formation reaction product, pressure 0.2MPaG, temperature 550 DEG C.Reaction product is divided into 2 strands after being drawn by the 2nd reactor---and reaction product 1, reaction product 2, throughput ratio is 0.7.Reaction product 1 and methanol of reaction charging 1 carry out heat exchange in the 2nd heat exchange unit, and methanol feeding 1 is heated to 250 DEG C.Reaction product 2 is divided into reaction product 3, reaction product 4, and throughput ratio is 1.1, carries out heat exchange respectively with the circulation gas from recycle gas compressor in the 3rd heat exchange unit, the 4th heat exchange unit.Reaction product 1 after heat exchange, reaction product 3, reaction product 4 are converged, and are cooled to 45 DEG C through wet type air cooling, enter the three phase separation that three phase separation tank carries out vapour, 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 0.2 or 0.3MPaG.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 intermediate product from the 1st reactor with the 1st reactor feed.

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.Component 1 enters recycle gas compressor, is pressurized to 0.32MPaG.The circulation gas leaving recycle gas compressor is divided into 3 strands---and circulation gas 1, circulation gas 2, circulation 3, throughput ratio is 1:0.9:1.1.Circulation gas 1 and reaction product 3 carry out heat exchange in the 3rd heat exchange unit, and circulation gas is heated to 480 DEG C.Circulation gas 2 and reaction product 4 carry out heat exchange in the 4th heat exchange unit, and circulation gas is heated to 490 DEG C.The raw material 1 in circulation gas 1 and the 1st heat exchange unit heat-absorbing medium exit converges, jointly as the reaction feed of the 1st reactor; The intermediates that circulation gas 2 and the 1st heat exchange unit exothermic medium export and raw material 2 converge, jointly as the reaction feed of the 2nd 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: 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 cools through the array mode of dry type air cooling, water-cooled, and temperature is down to 40 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, mixes with the 1st reactor product intermediates and raw material 2, participates in reaction as the 2nd reactor reaction charging, and the 5th heat exchange unit thermal source is outer for 1.2MPaG steam.

Circulation gas 3 enters the cold heat exchange unit of ammonia, and temperature is down to-13 ~ 30 DEG C, such as to 7 DEG C.The cold heat exchange unit of ammonia comprises gas-liquid separation device, and circulation gas 3 is condensed into two-phase after gas-liquid separation: liquid phase circulation gas 3 and gas phase 3.Liquid phase circulation gas 3 is pressurized to 0.5 ~ 1.8MPaG through pump, such as 1.35MPaG, enters de-liquefied gas tower.De-liquefied gas tower operating parameters is as follows: tower top pressure: 1.3MPaG; Tower reactor pressure: 1.35MPaG.Through de-liquefied gas tower fractionation, in liquid phase circulation gas 3, below C4 component is discharged by tower top, and C5 liquid-phase product is discharged by the bottom of tower, mixes with C5 liquid-phase product at the bottom of depentanizer, participates in reaction as the 2nd reactor reaction charging.De-liquefied gas column overhead gas phase cools through wet type air cooling mode, and temperature is down to 32 DEG C, enters de-liquefied gas return tank of top of the tower.C1 ~ C2 gaseous component is discharged by tank deck, and C3-C4 liquid phase is through de-liquefied gas trim the top of column pump supercharging, and part backflow returns de-liquefied gas column overhead; Another part C3 ~ C4 liquid-phase product is as the extraction of liquefied gas product.

The gas phase 2 that C1 ~ C4 gaseous component that depentanize return tank of top of the tower is discharged and de-liquefied gas return tank of top of the tower C1-C2 gaseous component of discharging and three phase separation tank is separated and the gas phase 3 that the cold heat exchange unit of ammonia is discharged enter 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 gas phase rising throughput ratio are 20.The temperature on this absorption tower is normal temperature, 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 three (see Fig. 3): containing the 1st, the 2nd, the 3 three reactor, separating step adopts depentanizer and the fractionation of de-liquefied gas tower, circulation gas 3 is first through residue gas compressor supercharging before entering de-liquefied gas tower, and methanol feedstock divides 3 stocks not enter three reactors.

Fresh methanol charging 1 pumping outside battery limit (BL), boosts to 0.6MPaG, temperature 30 DEG C.Fresh methanol charging 1 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 8:2:1.Raw material 1 enters the 1st reactor, carries out radially moving bed contact reacts with the highly active catalyzer from revivifier, and liquid hourly space velocity is 2.5h ^-1, generate intermediates 1 (i.e. the reaction product of the 1st reactor): pressure 0.56MPaG, temperature 440 DEG C.Intermediates 1 enter the 1st heat exchange unit after leaving the 1st reactor, and carry out preheating as thermal source to the 1st reactor feed, methanol feeding 1 is heated to 400 DEG C, and intermediates 1 are cooled to 420 DEG C.Be mixed into the 2nd reactor with raw material 2 after intermediates 1 leave the 1st heat exchange unit, carry out radially moving bed contact reacts with the catalyzer from the 1st reactor, liquid hourly space velocity is 2.5h ^-1, generate intermediates 2 (i.e. the reaction product of the 2nd reactor), pressure 0.54MPaG, temperature 480 DEG C.Intermediates 2 are drawn afterwards by the 2nd reactor and raw material 3 is mixed into the 3rd reactor, and carry out radially moving bed contact reacts with the catalyzer from the 2nd reactor, liquid hourly space velocity is 2.5h ^-1, formation reaction product, pressure 0.52MPaG, temperature 520 DEG C.

Reaction product is divided into 2 strands---and reaction product 1, reaction product 2, throughput ratio is 1.1.Reaction product 1 and methanol of reaction charging 1 carry out heat exchange in the 2nd heat exchange unit, and methanol feeding 1 is heated to 128 DEG C.Reaction product 2 is divided into reaction product 3, reaction product 4, reaction product 5, and throughput ratio is 4:6:2, carries out heat exchange respectively with from the circulation gas of recycle gas compressor and de-liquefied gas bottom product in the 3rd heat exchange unit, the 4th 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 are cooled to 40 DEG C through wet type air cooling, 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 0.9MPaG.This revivifier is conventional regeneration device.High activated catalyst after revivifier regeneration is promoted to the 1st reactor head, moving bed radial contact reacts is carried out with the 1st reactor feed, enter the 2nd reactor again, moving bed radial contact reacts is carried out with the intermediates from the 1st reactor, enter the 3rd reactor again, carry out moving bed radial contact reacts 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 11.0.Gas phase 1 enters recycle gas compressor, is pressurized to 0.7MPaG.The circulation gas leaving recycle gas compressor is divided into 3 strands---and circulation gas 1, circulation gas 2, circulation gas 3, throughput ratio is 0.8:1.0:1.2.Circulation gas 1 and reaction product 3 carry out heat exchange in the 3rd heat exchange unit, and circulation gas is heated to 400 DEG C.Circulation gas 2 and reaction product 4 carry out heat exchange in the 4th heat exchange unit, and circulation gas is heated to 480 DEG C.Circulation gas 1 mixes, jointly as the reaction feed of the 1st reactor with the raw material 1 after intensification; Intermediates 2 and the raw material 3 of circulation gas 2 and the 2nd reactor outlet mix, jointly as the reaction feed of the 3rd 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: 0.8MPaG; Tower reactor pressure: 0.85MPaG.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 cools through the array mode of dry type air cooling, water-cooled, and temperature is down to 80 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 200 DEG C after the 5th heat exchange unit and reaction product 5 heat exchange, mixes with the 2nd reactor product intermediates 2 and raw material 3, participates in reaction as the 3rd reactor reaction charging.

Circulation gas 3 enters residue gas compressor and is pressurized to 0.5 ~ 1.8MPaG, such as 1.41MPaG, then enters de-liquefied gas tower.De-liquefied gas tower operating parameters is as follows: tower top pressure: 1.4MPaG; Tower reactor pressure: 1.45MPaG.Through de-liquefied gas tower fractionation, in circulation gas 3, below C4 component is discharged by tower top, and C5 liquid-phase product is discharged by the bottom of tower, mixes with liquid-phase product at the bottom of depentanizer, returns the 3rd reactor feed.De-liquefied gas column overhead gas phase cools through wet type air cooling mode, and temperature is down to 34 DEG C, enters de-liquefied gas return tank of top of the tower.C1 ~ C2 gaseous component is discharged by tank deck, and C3 ~ C4 liquid phase is through de-liquefied gas trim the top of column pump supercharging, and part backflow returns de-liquefied gas column overhead; Another part C3 ~ C4 liquid-phase product is as the extraction of liquefied gas product.

The gas phase 2 that C1 ~ C4 gaseous component that depentanize return tank of top of the tower is discharged and de-liquefied gas return tank of top of the tower C1 ~ C2 gaseous component of discharging and three phase separation tank is separated 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 gas phase rising throughput ratio are 9.2.The temperature on this absorption tower is normal temperature, 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 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 four (see Fig. 4): containing the 1st, the 2 two reactor, separating step adopts depentanizer and the fractionation of de-liquefied gas tower.

Fresh methanol charging 1 pumping outside battery limit (BL), boosts to 0.8MPaG, temperature 25 DEG C.First fresh methanol charging 1 enters the 1st reactor (being equivalent to most top reactor) after heating up with reaction product heat exchange, and carry out radially moving bed contact reacts with the highly active catalyzer from revivifier, liquid hourly space velocity is 1.5h ^-1, generate intermediates (i.e. the reaction product of the 1st reactor), pressure 0.76MPaG, temperature 410 DEG C.Intermediates enter the 1st heat exchange unit after leaving the 1st reactor, and carry out preheating as thermal source to the 1st reactor feed, methanol feeding 1 is heated to 370 DEG C, and intermediates are cooled to 385 DEG C.Intermediates enter the 2nd reactor (being equivalent to least significant end reactor) after leaving the 1st heat exchange unit, carry out radially moving bed contact reacts with the catalyzer from the 1st reactor, and liquid hourly space velocity is 1.5h ^-1, formation reaction product, pressure 0.74MPaG, temperature 460 DEG C.Reaction product is divided into 2 strands after being drawn by the 2nd reactor---and reaction product 1, reaction product 2, throughput ratio is 0.74.Reaction product 1 and methanol feeding 1 carry out heat exchange in the 2nd heat exchange unit, and methanol feeding 1 is heated to 150 DEG C.Reaction product 2 is divided into reaction product 3, reaction product 4, and throughput ratio is 0.88, carries out heat exchange respectively with the circulation gas from recycle gas compressor in the 3rd heat exchange unit, the 4th heat exchange unit.Reaction product 1 after heat exchange, reaction product 3, reaction product 4 are converged, and are cooled to 40 DEG C through wet type air cooling, 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 580 DEG C, regeneration pressure 0.85MPaG.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 intermediate product from the 1st reactor with the 1st reactor feed.

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 17.0.Gas phase 1 enters recycle gas compressor, is pressurized to 0.86MPaG.The circulation gas leaving recycle gas compressor is divided into 3 strands---and circulation gas 1, circulation gas 2, circulation 3, throughput ratio is 0.92:1.0:0.98.Circulation gas 1 and reaction product 3 carry out heat exchange in the 3rd heat exchange unit, and circulation gas is heated to 370 DEG C.Circulation gas 2 and reaction product 4 carry out heat exchange in the 4th heat exchange unit, and circulation gas is heated to 385 DEG C.The fresh methanol charging 1 in circulation gas 1 and the 1st heat exchange unit heat-absorbing medium exit converges, jointly as the reaction feed of the 1st reactor; The intermediates that circulation gas 2 and the 1st heat exchange unit exothermic medium export converge, jointly as the reaction feed of the 2nd 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: 0.7MPaG; Tower reactor pressure: 0.75MPaG.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 cools through the array mode of dry type air cooling, water-cooled, and temperature is down to 110 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 160 DEG C through the 5th heat exchange unit, mixes with the 1st reactor product intermediates, participates in reaction as the 2nd reactor reaction charging, and the 5th heat exchange unit thermal source is outer for 1.2MPaG steam.

Circulation gas 3 enters de-liquefied gas tower.De-liquefied gas tower operating parameters is as follows: tower top pressure: 0.3 or 0.4MPaG; Tower reactor pressure: 0.35 or 0.45MPaG.Through de-liquefied gas tower fractionation, in circulation gas 3, below C4 component is discharged by tower top, and C5 liquid-phase product is discharged by the bottom of tower, mixes with C5 liquid-phase product at the bottom of depentanizer, participates in reaction as the 2nd reactor reaction charging.De-liquefied gas column overhead gas phase cools through wet type air cooling mode, and temperature is down to 32 DEG C, enters de-liquefied gas return tank of top of the tower.Wherein C4 liquid phase component is sunken at the bottom of tank, C1 ~ C3 gaseous component is discharged by the tank deck of return tank of top of the tower, 1.3MPaG is pressurized to through circulation residue gas compressor, again through wet type air cooling, the cooling of water-cooled array mode, temperature is down to 32 DEG C, enters de-liquefied gas tower top knockout drum, wherein C1 ~ C2 gaseous component is discharged by the tank deck of tower top knockout drum, enter post-processing step, C3 liquid-phase reflux, to de-liquefied gas return tank of top of the tower, converges with C4 liquid phase component.C3 ~ C4 liquid phase is through de-liquefied gas trim the top of column pump supercharging, and part backflow returns de-liquefied gas column overhead; Another part C3 ~ C4 liquid-phase product is as the extraction of liquefied gas product.

The C1-C2 gaseous component that the C1-C4 gaseous component of being discharged by depentanize return tank of top of the tower top and de-liquefied gas tower top knockout drum are discharged and the gas phase 2 that three phase separation tank is separated enter 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 8.The temperature on this absorption tower is normal temperature, working pressure 1.0MPaG.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.

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

Table 3 liquefied gas product forms

Table 4 non-condensable gas forms

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. the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of a co-production of liquefied gas, comprise hydrocarbon synthesis and separating 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 most top reactor after heating up, its reaction product enters a reactor thereafter as reaction raw materials, the rest may be inferred, until the reaction product of penultimate reactor enters least significant end reactor as reaction raw materials, described reactor is radially moving bed reactor, the reaction product of multiple reactor successively carries out heat exchange with methanol feedstock as exothermic medium, cascade raising temperature is carried out to methanol feedstock, the plurality of reactor at least comprises most top reactor and least significant end reactor, the reaction product of least significant end reactor is divided into multiply, wherein at least one for heating up to methanol feedstock, entered the temperature of charge of most top reactor according to the temperature of the reaction product of least significant end reactor by the flow control changing this burst of reaction product, and then control the temperature of reaction of most top reactor, described separating step adopts gas-oil-water three-phase separating device to converge each stock and the reaction product of cooled least significant end reactor carries out three phase separation, is separated the aqueous portion obtained and sends into oil-contained waste water treatment device, be separated the oil phase part obtained and be distributed into depentanizer separation and Extraction C6 ~ C10 aromatic hydrocarbon product, be separated the most of gas phase obtained and be used as circulation gas through recycle gas compressor compression, wherein return the reactor of hydrocarbon synthesis step after most of circulation gas intensification as reaction raw materials, small portion circulation gas enters de-liquefied gas tower fractionation extraction C3 ~ C4 and draws as liquefied gas product.

2. the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas as claimed in claim 1, it is characterized in that described most of circulation gas enters most top reactor and least significant end reactor respectively as reaction raw materials after being divided into two strands to heat up separately, by the reaction depth regulating these two strands of circulation gas flows independently to regulate most top reactor and least significant end reactor online.

3. the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas as claimed in claim 2, it is characterized in that before each stock of the reaction product of least significant end reactor converges, also have two strands in multiply reaction product and by heat exchange, per share circulation gas is heated up respectively as exothermic medium.

4. the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas as claimed in claim 3, the method that it is characterized in that adopting de-liquefied gas tower fractionation to extract C3 ~ C4 be following any one: (1): C5 liquid phase is discharged by the bottom of tower, returns least significant end reactor after heating up as reaction raw materials, top gaseous phase C1 ~ C4 temperature is down between the boiling point of C2 and C3 under logistics current pressure, enter return tank of top of the tower, wherein C1 ~ C2 gaseous component is discharged by the tank deck of return tank of top of the tower, enters post-processing step, C3 ~ C4 liquid phase is through the supercharging of trim the top of column pump, and part backflow returns tower top, another part is as the extraction of liquefied gas product, (2) make C5 liquid phase discharge by the bottom of tower, after heating up as reaction raw materials, return most top reactor, top gaseous phase C1 ~ C4 temperature is down between the boiling point of C3 and C4 under logistics current pressure, enter return tank of top of the tower, wherein C1 ~ C3 gaseous component is discharged by the tank deck of return tank of top of the tower, be cooled between the boiling point of C2 and C3 under logistics current pressure again through the supercharging of circulation residue gas compressor, enter de-liquefied gas tower top knockout drum, C1 ~ C2 gaseous component after separation is discharged by the tank deck taking off liquefied gas tower top knockout drum, enter post-processing step, C3 liquid-phase reflux after separation is to de-liquefied gas return tank of top of the tower, converge with C4 liquid phase component, C3 ~ C4 liquid phase is through the supercharging of trim the top of column pump, part backflow returns tower top, another part is as the extraction of liquefied gas product.

5. the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas as claimed in claim 4, is characterized in that the process of employing depentanizer separation and Extraction aromatic hydrocarbon product is: make 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, wherein C1 ~ C4 gaseous component is discharged by the tank deck of return tank of top of the tower, enter post-processing step, C5 liquid phase is through the supercharging of trim the top of column pump, and part backflow returns tower top; Another part returns least significant end reactor after heating up as reaction raw materials.

6. the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas as claimed in claim 5, it is characterized in that described post-processing step adopts methyl alcohol to treat treated substance and carries 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.

7. as the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of the co-production of liquefied gas in claim 1-6 as described in any one, it is characterized in that described small portion circulation gas to enter before de-liquefied gas tower first through residue gas compressor supercharging or first through the cold heat exchange unit cooling of ammonia again through pump supercharging.

8., as the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of the co-production of liquefied gas in claim 1-7 as described in any one, it 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 0.25 ~ 1.75MPaG, temperature is 320 ~ 520 DEG C, the pressure of least significant end reactor is 0.2 ~ 1.73MPaG, temperature is 370 ~ 550 DEG C, after often in adjacent two reactors, the top pressure of a reactor is not higher than the top pressure of last reactor, and all not higher than the top pressure of most top reactor, the minimal pressure of a rear reactor is not higher than the minimal pressure of last reactor, and all not higher than the minimal pressure of most top reactor, 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 270 ~ 490 DEG C, methanol feedstock is successively warming up to 98 ~ 250 DEG C and 250 ~ 480 DEG C, liquid at the bottom of absorption tower first mixes again with methanol feedstock together cascade raising temperature with methanol feedstock, 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 de-liquefied gas tower is 0.3 ~ 1.6MPaG, tower reactor pressure is 0.35 ~ 1.65MPaG, the C5 returning least significant end reactor is warming up to 150 ~ 250 DEG C by heat exchange, the reaction product that thermal source is reactor described in hydrocarbon synthesis step or outer supplying heat source.

9. the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of co-production of liquefied gas as claimed in claim 8, it is characterized in that being provided with gas-liquid separation device in the cold heat exchange unit of described ammonia, the cold heat exchange unit of described ammonia is separated into more than C3 liquid composition and C1 ~ C2 gaseous component by after described small portion circulation gas cooling, enter de-liquefied gas tower after the supercharging of more than C3 liquid composition, C1 ~ C2 gaseous component enters post-processing step; Described small portion circulation gas first through residue gas compressor supercharging, is pressurized to 0.5 ~ 1.8MPaG; First through the cooling of ammonia cold heat exchange unit again through pump supercharging, be cooled to-13 ~ 30 DEG C, be pressurized to 0.5 ~ 1.8MPaG.

10. as the High Efficiency Thermal integrated-type moving-bed aromatization of methanol method of the co-production of liquefied gas 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 most top reactor, other each stocks do not enter other reactors, and the methanol feedstock accounting entering most top reactor is greater than the methanol feedstock entering other each reactors.