Since its discovery in the 1970's, zeolite catalyzed conversion of methanol to hydrocarbons has become increasingly important in the chemical industry and several variations of the process have been commercialized including MTO (methanol-to-olefins), MTP (methanol-to-propylene), and MTG (methanol-to-gasoline). Herein, the focus is on the MTA (methanol-to-aromatics) process, in which methanol is converted over a metal/oxide containing zeolite to a mixture of hydrocarbons with a high content of aromatic compounds. In the MTA process accumulation of hydrogen in the MTA loop may inhibit dehydro- genation reactions taking place over the MTA catalyst resulting in lower yield of aro- matics. Known methods for removal of hydrogen, include removal of H2 from a recycle, e.g. by means of a perm-selective membrane and/or by selective oxidation in the recycle stream. However, there is still a need for alternatives for reducing the H2 concentra- tion in the MTA loop.

The present invention relates to a process for converting a feed stream comprising oxygenates to a product stream rich in aromatics, said process comprising the steps of: providing a synthesis gas from a process front end based on natural gas and/or ethane

converting said synthesis gas to a feed stream comprising oxygenates providing the feed stream comprising oxygenates to a reactor wherein an oxygenate to aromatics conversion process takes place,

separating the conversion effluent from the reactor into at least a liquid hydro- carbon product phase, a gas phase, and an aqueous condensate,

adding a co-feed to the feed stream/MTA loop, wherein

the co-feed is obtained from one or more off- streams from the process frontend, and

optionally recycling at least part of the gas phase,

In the present process, aromatics are produced from natural gas and/or ethane via synthesis of methanol and/or dimethyl ether (DME) from a synthesis gas obtained from the natural gas and/or ethane. The liquid hydrocarbon product obtained by this process may be referred to as "reformate". In particular the applicant has observed that adding a co-feed of a stream comprising C02 and/or a stream comprising N2 to the MTA loop leads to higher yields of aromatics. The yield of aromatics is defined as the amount (moles) of carbon in the produced aromatics divided by the amount (moles) of carbon in the oxygenates in the feed stream to the MTA reactor.

Introduction of a stream comprising C02 and/or a stream comprising N2 as co-feed to the MTA reactor may be accompanied by a higher purge flow, which results in depletion of H2 from the MTA loop. Lower H2 concentration in the MTA reactor leads to higher dehydrogenation activity of the MTA catalysts, which ultimately results in a higher yield of aromatics. C02 and/or N2 passes through the MTA reactor substantially unconverted, even though C02 may participate in the reverse water-gas-shift reaction. Thus, the role of the co-feed is simply to act like an inert gas sweeping H2 from the loop. Furthermore, the use of C02 as co-feed has the advantage that C02 has a high heat capacity and it is therefore a good heat sink for the MTA reaction system.

The higher purge flow obtained from adding C02 and/or N2 co-feed may result in loss of intermediates/precursors for aromatics from the MTA loop (for instance olefins) as well as aromatics themselves (for instance benzene) through the purge, but this effect is only significant at high co-feed flows.

The natural gas reforming section may comprise various reforming steps. Preferably the process front end comprises an air separation unit (ASU) step, one or more reforming steps, and/or a C02 removal step.

Synthesis gas is a mixture of predominately hydrogen and carbon monoxide (normally more than 80 vol. %). Synthesis gas may also contain other components in smaller concentrations such as carbon dioxide, methane, nitrogen and argon. The synthesis gas may for example be produced by technologies such as: a) Thermal Partial Oxidation without catalyst (POX)

b) Catalytic Partial Oxidation (CPO)

c) Autothermal Reforming (ATR) In addition or in alternative to the technologies a - c described above other reforming techniques may also be used:

Two step reforming. In this case natural gas is fed to a steam reformer (SMR). The effluent from the steam reformer is sent to an Autothermal Reformer.

Heat Exchanger Reforming in series with the Autothermal Reformer. In this case part or all of the natural gas feed is sent to a heat exchange reformer. The effluent from the heat exchange reformer is mixed with any remaining feed and sent to the ATR. Part or all of the ATR effluent is sent to the heat exchange reformer and cooled while supplying heat to the endothermic steam reforming reaction.

Heat Exchange Reforming in parallel with the Autothermal Reformer. In this case the natural gas feed is divided into two parts. The largest part of the feed is sent to the ATR. The remaining feed is sent to the heat exchange reformer. The effluent from the heat exchange reformer and part or all of the effluent from the ATR are mixed to provide a combined synthesis gas stream. The combined synthesis gas stream is cooled in the heat exchange reformer supplying heat for the endothermic steam reforming reaction.

In all of the above cases part or all of the natural gas and/or ethane may optionally be pre-reformed before being fed to the various reactors.

The term "pre-reforming" and "pre-reformer" as used herein is defined as a steam reforming process and steam reformer by which higher hydrocarbons (with 2 or more carbon atoms) are converted to a mixture of methane, carbon oxides and hydrogen at temperatures in the range 375-650°C, more specifically 400-600°C, preferably adiabati- cally in a fixed bed of catalyst comprising nickel, and its main purpose is to remove hydrocarbons higher than methane.

In another embodiment, each individual stream in the form of first hydrocarbon feedstock, or second hydrocarbon feedstock, or both, are subjected to pre-reforming prior to passing through autothermal reforming or heat exchange reforming. During pre-reforming, most or all of the higher hydrocarbons (hydrocarbon compounds with 2 or more carbon atoms) are converted according to the following reactions:

CnHm + nH20→ (½m+n)H2 + nCO (1 )

3H2 + CO <→ CH4 + H20 (2)

CO + H20 ^ H2 + C02 (3) Reactions (2) and (3) are normally close to equilibrium at the outlet of the pre-reformer.

Autothermal reforming (ATR) is described widely in the open literature. Typically, the autothermal reformer comprises a burner, a combustion chamber, and catalyst arranged in a fixed bed all of which are contained in a refractory lined pressure shell. Au- tothermal reforming is for example described in Chapter 4 in "Studies in Surface Science and Catalysis", Vol. 152 (2004) edited by Andre Steynberg and Mark Dry.

In the ATR, hydrocarbon feedstock, oxidant gas, and in some cases steam is added. "Raw synthesis gas" is formed by a combination of partial oxidation and steam reform- ing in the autothermal reformer. By the term "raw synthesis gas" is meant a gas mixture obtained from the synthesis gas production unit, from which C02 has not been removed. After removal of C02 (and optionally addition of H2), the gas is termed "synthesis gas", which may enter the oxygenate synthesis reactor. If no C02 removal is needed, the "raw synthesis gas" can be used directly as "synthesis gas" in the oxygen- ate synthesis reactor.

By the term "oxidant gas" is meant a stream comprising oxygen, preferably more than 75 vol%, more preferably more than 85 vol% oxygen and most preferably more than 95% oxygen. Examples of oxidant gas are oxygen, mixture of oxygen and steam, mix- tures of oxygen, steam, and argon, and oxygen enriched air.

When oxidant gas is produced, a gas which is rich in N2 is formed as a byproduct. This N2 rich gas may also contain argon, and may be used as co-feed to the MTA reactor. The temperature of the synthesis gas leaving is between 900 and 1 100°C, or 950 and 1 100°C, typically between 1000 and 1075°C. This hot effluent synthesis gas leaving the autothermal reformer comprises carbon monoxide, hydrogen, carbon dioxide, steam, residual methane, and various other components including nitrogen and argon.

Fuel gas is used in the fired heaters to preheat the feedstock upstream the reforming step. Typically a second preheater is also present producing steam to cover the process energy demands. The raw synthesis gas withdrawn from the one or more reforming steps may be too lean in hydrogen to be a suitable feed for the subsequent methanol synthesis. The desired stoichiometry for the methanol synthesis is often characterised by the so called module, M: M = (XH2— Xco2) / (Xco + Xco2)

XH2, XCO, and Xco2 are the mole fractions of hydrogen, carbon monoxide, and carbon dioxide respectively. The desired value of M will typically be around 2 or slightly above. The raw synthesis gas withdrawn from the synthesis gas production step may have a module value less than 2, such as less than 1 ,9, or even less than 1 .8. In order to modify the module, M, several means are available. One way is to increase the concentration of hydrogen. Another option is to remove part of the carbon dioxide from the raw synthesis gas. This can be done by various means such as washing with methanol, us- ing an amine wash step, rectisol, selexol or similar commercial processes.

When C02 is removed from the raw synthesis gas, a C02 rich stream is produced. This stream may be used as co-feed for the MTA reactor. C02 may to a limited extent react with H2 by the reverse shift reaction (3) depending upon process conditions. This reduces the H2 concentration in the reactor and may lead to higher yield of aromatics.

The synthesis gas leaving the synthesis gas production unit is fed to a methanol or DME synthesis unit or section to produce methanol. The synthesised methanol and/or DME may contain minor amounts of by-products and dissolved gases. Hence, the raw methanol may optionally be purified for example by distillation in one or more columns to obtain the desired purity and quality.

In the MTA process oxygenates are converted to hydrocarbons. The oxygenates may preferably be methanol and/or dimethylether (DME). The MTA conversion process may be carried out at a pressure of 5-60 bar, preferably 10-40 bar, at a temperature of 300- 500°C, preferably 300-480°C, and a weight hourly space velocity (kg methanol and/or DME feed per kg of catalyst per hour) between 0.1 and 10, preferably 0.5-3. In case the MTA reactor is an adiabatic fixed bed reactor, the difference between the inlet temperature and the outlet temperature is preferably between 30 and 150°C, more preferably between 50 and 130°C. Preferably the inlet temperature to the reactor is between 320°C and 380°C and the outlet temperature is between 380°C and 480°C. The present process may provide a product particularly suited as feedstock for downstream aromatics processing, e.g. for producing para-xylene. The liquid hydrocarbon product may comprise various aromatics such as benzene, toluene, xylenes, ethylben- zene, and heavier aromatic compounds with 9 or more carbon atoms, as well as n-par- affins, isoparafins, olefins, and naphthenes. Most of the hydrocarbons present in the liquid hydrocarbon stream may be components with 4 or more carbon atoms, but the stream may also comprise lighter hydrocarbons in low concentrations as well as small amounts of dissolved gases such as CO2, CO, and H2.

The first gas stream from the separator comprises mainly light hydrocarbons such as methane, ethane, propane, butanes, ethylene, propylene and/or butylenes, as well as carbon oxides, carbon dioxide, and hydrogen. If large amounts of N2 and/or C02 are used as co-feed, these species will be found in large amounts in the gas stream leaving the separator. The aqueous condensate comprises mainly water, but may also comprise small amounts of various oxygenates including methanol, other alcohols, aldehydes, and ketones as well as dissolved gases. A typical industrial unit will produce 100 - 2500 kiloton/year of liquid hydrocarbon product (reformate). If the oxygenate synthesis section is off-line, the MTA process may be performed using methanol and/or DME from another source. In order to achieve a desired selectivity to aromatics, an MTA catalyst preferably comprises a zeolite or zeotype as well as a metal/oxide function. The zeolite/zeotype is responsible for conversion of oxygenates to hydrocarbons, while the metal/oxide function is responsible for dehydrogenation of intermediate hydrocarbons, e.g. dehydrogenation of naphthenes to aromatics and/or dehydrogenation of paraffins to olefins. The combi- nation of a zeolite function and a dehydrogenation function is essential for achieving a high yield of aromatics in the MTA process.

Different zeolite/zeotypes may be employed, including ZSM-5, ZSM-1 1 , ZSM-23, ZSM- 48, SAPO-34, however ZSM-5 may be preferred, since it has a suitable size selectivity to the desired methylated monocyclic aromatic species as well as a relatively low coking rate. The metal component of the MTA catalyst may in advantageous embodiments be chosen from Zn, Ga, In, Ge, Ni, Mo, P, Ag, Sn, Pd and Pt or combinations thereof. Zn may be preferred over the other metals. Thus, a catalyst comprising Zn/ZSM-5 may be a particularly preferred catalyst system for the MTA process. Furthermore, the MTA catalyst may comprise phosphorus, which leads to better hydrothermal stability of the catalyst and thus longer ultimate catalyst lifetime.

It may be preferred to use a binder material in order to shape the catalyst. This binder material may be a normally employed binder material such as AI203, MgAI204, Si02, Zr02, Ti02, MgO or mixtures thereof. AI203 may be preferred. If AI203 is used as binder, Zn may be present in the catalyst as ZnAI204. Similarly, if AI203 is used as binder P may be present in the catalyst as AIP04.

An MTA catalyst may comprise 0.2 - 15 wt% Zn, or more preferably 3 - 15 wt% Zn or even more preferably 5 - 15 wt% Zn. Furthermore an MTA catalyst may comprise 0 -

10 wt% P, or more preferably 0.1 - 8 wt% P or even more preferably 0.5 - 5 wt% P.

The MTA process may be carried out in one or more fixed bed and/or fluid bed reactors. The present invention enables the control and optimization of the MTA process and yield by controlling the content and ratio of various streams, recycle and added co- feed. The molar ratio of co-feed to methanol and/or DME in the feed is preferably 0.01 - 10, more preferably 0.02 - 1 or most preferably 0.05 - 0.5. If it is too low, the desired effect will not be significant. If it is too high, loss of intermediates through the purge may lead to lower yield of aromatics. Traditionally, a large part (e.g. more than 99%) of the gas from the separator is recycled to the reactor. However, the applicants have found that with the addition of the co- feed to the feed stream, a higher content of aromatics in the liquid hydrocarbon stream may be obtained by reducing the ratio of the gas phase from the separation which is recycled to the MTA reactor. Hence, along with the addition of a co-feed, the molar ratio of the amount of gas being purged from the MTA loop to the total amount of gas recovered from the separation system is preferably more than 1 %, such as more than 5% or more than 20% or more even more than 80%.

In another advantageous embodiment of the invention, the flow of recycle gas form the separator in the MTA loop to the inlet of the MTA reactor is very low, preferably 0, resulting in a low content of H2 at the inlet to the MTA reactor. The low concentration of H2 at the inlet of the MTA reactor may lead to high dehydrogenation activity of the MTA catalyst, and thus higher yield of aromatics and/or olefins. In low recycle embodiments, the ratio of recycle to co-feed may preferable be below 1 , or more preferably below 0.5 or even more preferable below 0.1 or most preferably 0. Embodiment of the present invention with a low recycle/co-feed ratio may be advantageous if suitable co-feed gas is available in sufficient amounts to control the temperature in the MTA reactor. If there is no (or limited) gas recycle from the separator in the MTA loop, the temperature increase in the MTA reactor may be high, if too little co-feed is added.

The concentration of H2 in the feed stream to the MTA reactor may preferably be as low as possible in order to have a high dehydrogenation activity of the MTA catalyst and thus obtain a high yield of aromatics. The concentration of H2 in the feed stream to the MTA reactor may preferably be below 2 mol% or more preferably below 1 mol% or most preferably below 0.5 mol%.

However, it may be advantageous to maintain a small fraction of H2 in the feed stream to the MTA reactor in order to suppress coking.

The process conditions (e.g. temperature) of the separator in the MTA loop may be varied to control the composition of the resulting gas and liquid streams. If no or very little recycle of the gas from the separator in the MTA loop is applied, it may be advan- tageous to use a lower temperature in the separator in order to get as much as possible of the C2+ hydrocarbons into the liquid stream. The temperature in the separator may be below 60°C or below 50°C, or more preferably below 40°C or even more preferably below 30°C. This may ultimately lead to a more efficient overall process, since part of the liquid stream may be recycle to the front end or the MTA loop.

The applicant has furthermore found that it is desirable if the content of hydrogen in the co-feed is low. A low concentration of hydrogen in the co-feed will result in a higher content of aromatics in the liquid hydrocarbon stream. Hence, the mole percentage of hydrogen in the co-feed stream should be less than 5%, such as less than 2%, more preferably less than 1 %, or most preferably less than 0.1 %. If it is too high, the desired effect will not be significant. Should preferably be as low as possible (preferably 0).

The applicant has furthermore discovered that the mole percentage of hydrogen in the co-feed preferably is less than the mole percentage of hydrogen in the recycle stream. This will result in a higher yield of aromatics. Hence, the mole fraction of molecular hydrogen in the co-feed should be lower than the mole fraction of molecular hydrogen in the recycle stream. Preferably the ratio of molar percentage of H2 in the co-feed to the molar percentage of H2 in the recycle is below 0.8, such as below 0.6, below 0.4 or more preferably below 0.2 or below 0.1 . If it is too high, the desired effect will not be significant. Preferably as low as possible (preferably 0).

The co-feed may further comprise methane, ethane, propane, LPG or higher hydrocarbons, either from an external source or from one or more recycle streams. Preferably the molar ratio of gas recycle from the separator in the MTA loop to the sum of co-feed and MeOH and/or DME in feed is 0 - 15, preferably 1 - 10, or more preferably 2 - 6. The molar ratio of gas recycle from the separator in the MTA loop to co-feed is preferably 0 - 100, more preferably 10 - 100 or most preferably 20 - 50.

If a relatively high flow of co-feed is added, high concentration of co-feed C02 and N2 are found in the purge. The purge may also contain methane, ethane, propane, eth- ylene, propylene, CO, H2, steam, C4+ hydrocarbons, and/or oxygenates. The purge gas may be used as fuel.

An overview of the present process is given in Figure 1. Stream 1 is a feed stream comprising natural gas and/or ethane entering the synthesis gas production unit (3). Stream 2 is a feed stream comprising steam, while stream 4 comprises air entering the Air Separation Unit, ASU (5). Stream 6 is a stream comprising oxygen or oxygen enriched air entering the synthesis gas production unit (3), while stream 7 is a stream comprising N2, which is split into a N2 purge (8) and a N2 co-feed stream (9) for the MTA reactor (19). Stream 10 is raw synthesis gas entering a C02 removal unit (1 1 ). A stream comprising C02 (14) from the C02 removal unit is split into a C02 purge (15) and a C02 co-feed stream (16) for the MTA reactor. Stream 12 is synthesis gas from the C02 removal unit (1 1 ) entering the oxygenate synthesis unit (13). Stream 17 is a stream comprising oxygenates such as methanol and/or DME, which is mixed with a co-feed stream comprising C02 (16) and/or a co-feed comprising N2 and/or a gas re- cycle stream from the MTA reactor effluent (26) to obtain the feed stream (18) for the MTA reactor (19). The effluent from the MTA reactor (20) enters a separator (21 ), where it is separated into an aqueous condensate (22), a liquid hydrocarbon product stream (23) and a first gas stream (24), which is split into a purge stream (25) and a recycle stream (26).