The conversion of methanol (MeOH) and/or dimethyl ether (DME) , into hydrocarbon fuels and chemicals is of signifi ^¬ cant industrial importance. Prominent examples comprise methanol-to-gasoline (MTG) and methanol-to-olefins (MTO) processes . In general, any alcohol or ether may be converted into hy ^¬ drocarbons by these processes, but methanol is the pre ^¬ ferred feedstock, because it may be produced in large scale and with high efficiency from any carbon-containing resource, such as coal, biomass, waste and natural gas. Prior to the conversion into hydrocarbons the methanol or alcohol feed may be converted, at least partially, into its ether analogue e.g. methanol to DME.

Similar processes have been disclosed for the conversion of alcohols into hydrocarbons, targeting different products such as aromatics.

The MTG process typically provides high-octane synthetic gasoline in excellent yields, close to about 90 per cent by weight of the total amount of hydrocarbons produced and with typical octane numbers close to 90, expressed in terms of (R+M) /2 (Research Octane Number (RON) + Motor Octane Number (MON)/2). The synthetic gasoline typically contains substantial amounts of aromatics which contribute signifi- cantly to the octane number. Even higher octane numbers may be obtained by increasing the aromatics content, e.g. by increasing the pressure or temperature in the gasoline syn ^¬ thesis reactor or by reducing the space velocity. Changing synthesis conditions to improve aromatics selectivity, and thereby octane, however, imposes a penalty in terms of a reduced gasoline yield, because the formation of aromatics is inherently accompanied by increased selectivity to par ^¬ affins, in particular, propane and butanes which only to a limited extent may be added to the gasoline product slate due to their relatively high vapour pressures. In other words, the existing choice is between making a low-aromat- ics product in high yields or a high-aromatics product in low yields.

In a first aspect of the present invention is provided a process and plant for producing a hydrocarbon product with enhanced aromatics selectivity without simultaneously in ^¬ creasing the selectivity to light paraffins such as propane and butanes. In a second aspect of the present invention is provided a process for the production of gasoline with improved yield and octane number.

In a third aspect of the present invention is provided a process for making a highly aromatic product.

These and other advantages are provided by a process for production of hydrocarbons comprising the steps of

- converting a feed stream comprising alcohols, ethers or mixtures hereof over a bifunctional catalyst comprising ze ^¬ olite, alumina binder and Zn, wherein the Zn is present at least partly as ΖΓ1ΑΙ ₂ Ο ₄ , active in dehydrogenation of hy ^¬ drocarbons, in a conversion step thereby obtaining a conversion effluent,

- separating said effluent to obtain an aqueous process condensate stream, a liquid hydrocarbon stream and a gase ^¬ ous stream,

- removing part of the hydrogen formed in the conversion step, and

recycling at least part of the gaseous and/or liquid hydro- carbon stream to the conversion step.

I.e. according to the present invention there is provided a process for making a hydrocarbon product in which hydrogen produced in the conversion step is removed or partially re- moved from the synthesis loop, either by purging part of the gaseous recycle stream, hereinafter referred to as pas ^¬ sive H ₂ rejection, or by removing, partially or completely, ¾ from one or more recycle streams, hereinafter referred to as active ¾ rejection, in order to provide an at least partly ¾ depleted recycle. In the process e.g. methanol is converted over a preferably acidic zeolite based catalyst, combined with a metal or metal oxide component active in the dehydrogenation of hydrocarbons, into a conversion ef ^¬ fluent comprising aromatics.

The conversion effluent is separated into a gaseous stream, an aqueous process condensate and a liquid hydrocarbon stream comprising a hydrocarbon product. By the present process and plant wherein hydrogen is re ^¬ moved it is thus possible to increase the aromatics selec ^¬ tivity and at the same time reduce the selectivity to light paraffins, such as propane and butanes.

When the process is carried out with a bifunctional cata ^¬ lyst for conversion of oxygenates and dehydrogenation of hydrocarbons, said catalyst comprising zeolite, alumina binder and Zn, wherein the Zn is present at least partly as ΖηΑΐ ₂ θ ₄ the aromatics yield is increased compared to a standard catalyst without Zn. Furthermore, the applicant has shown that the selectivity of the catalyst is increased with the spinelization degree of the catalyst. To increase the yield of aromatics, a bifunctional catalyst containing acidic zeolite sites as well as dehydrogenation sites e.g. metal or oxide is provided. This means that a stream comprising one or more oxygenates e.g. methanol may be converted in the presence of the catalyst into aromatics by the zeolite while dehydrogenation of hydro-carbons such as naphthenes, paraffins and/or isoparaffins , into olefins and/or aromatics also takes place.

In preferred embodiments the catalyst is optimized for con- version of oxygenates such as Methanol and/or DME into aro ^¬ matics (herein abbreviated MTA) .

The binder may be a pure alumina binder or an alumina-based binder further comprising mixtures of aluminum oxide and aluminum hydroxide and/or e.g. silica/alumina. The zeolite may for example be one of the commonly known zeolites used in MTA and MTG processes. For example, H-ZSM- 5 may be a preferred zeolite for the present catalyst due to its unique pore structure leading to favorable size se- lectivity as well as its relatively low coking rate. H- ZSM-5 may be particularly preferred in case of MTA pro ^¬ cesses.

Examples of Zn/ZSM-5 catalysts with low content of Zn such as 1 wt% Zn for MTA are known and it has been argued that higher Zn content is to be avoided in order to avoid metha ^¬ nol cracking to carbon oxides. However, the applicant has shown that a high Zn content in the catalyst may result in an improved aromatics yield in MTA processes compared to known catalysts. Thus, in several advantageous embodiments the total Zn content in the catalyst is 3 - 25 wt%, 5 - 20 wt%, 7 - 15 wt% or 8 - 13 wt%, such as more than 7 wt% Zn, more than 10 wt% Zn or 12 wt% or more Zn. Depending on the production process the Zn in the catalyst may be present in various concentrations in both binder and zeolite of the present catalyst. E.g. in some embodiments the Zn concentration is higher in the binder phase than in the zeolite phase which for example may be the case where the Zn is applied by impregnation.

A catalyst wherein Zn is present in both zeolite and alu ^¬ mina binder allows for industrial production by "simple" means such as by impregnation. For example, a bifunctional catalyst as herein described may be achieved by Zn impreg ^¬ nation of a "base catalyst" comprising an alumina binder and a zeolite such as ZSM-5. A preferred base catalyst com ^¬ prises 30-50 % binder and 50-70 % zeolite.

The impregnation may be carried out by contacting the zeo- lite or the zeolite and alumina binder with a Zn-containing solution. The solution may preferably be aqueous, but other solvents than water may be preferred as well. Impregnation may also be carried out by contacting the zeolite or the zeolite and alumina binder with a solid Zn compound, e.g., by mixing and/or grinding or other treatments to ensure intimate mixing of the components.

The Zn source may be any Zn-containing, organic and/or inorganic, compound. Preferred compounds comprise zinc ni- trate, zinc acetate and zinc oxide, hydroxide, carbonate or mixtures hereof.

In order to provide a functional catalyst, the impregnation will typically be followed by calcination or similar treat- ment ( s ) .

However, when a zeolite or an alumina/zeolite based cata ^¬ lyst is impregnated with Zn in order to obtain the desired amount of Zn in the zeolite, significant amounts of Zn may also be introduced into the binder, for example, as ZnO and/or ΖηΑΐ ₂ θ ₄ . Various ratios of ΖηΟ/ ΖηΑΐ ₂ θ ₄ may be

achieved depending on the treatment of the impregnated cat ^¬ alyst. The applicant has shown that in a desirable catalyst Zn in the alumina binder is present mainly as ΖηΑΐ ₂ θ ₄ . Defining the relative amount of zinc oxide, ZnO, in the binder phase as molar percentage of Zn present as ZnO relative to the total amount of Zn contained in the binder phase it may be desirable to have a catalyst where the amount of ZnO pre ^¬ sent in the binder phase as less than 50%, or preferably less than 10%, such as less than 5% or less than 2%, pref ^¬ erably less than 1%, such as 0.5% or less than 0.1% ZnO.

I.e. it may be preferred that the Zn in the binder has been fully spinelized, according to the reaction equation ZnO + AI ₂ O ₃ ΖηΑΐ ₂ θ ₄ , meaning that all or substantially all of the Zn in the binder is present as ΖηΑΐ ₂ θ ₄ .

Preferably a large part of the Zn in the alumina binder is present as ΖηΑΐ ₂ θ ₄ . Defining the relative amount of ΖηΑΐ ₂ θ ₄ in the binder phase as molar percentage of Zn present as

ΖηΑΐ ₂ θ ₄ relative to the total amount of Zn contained in the binder phase, in some embodiments 50 - 100% of the Zn in the binder is present as ΖηΑΐ ₂ θ ₄ , for example more than 60%, more than 70% or more than 80%. In some advantageous embodiments 85 - 100% of the Zn in the binder is present as ΖηΑΐ ₂ θ ₄ , such as more than 90% or more than 95%.

As shown by the applicant cracking of MeOH may be avoided with a high degree of spinelization, it may be preferred especially in case of high Zn content in the catalyst that more than 97% of the Zn in the binder is present as

ΖηΑΐ ₂ θ ₄ , such as more than 98%, more than 99%, more than 99,5% or more than 99,8% of the Zn in the binder is present as ΖηΑΐ ₂ θ ₄ . Optimal and practically achievable ΖηΑΐ ₂ θ ₄ con- tent ranges may be 95 - 100% in the binder is present as

ΖηΑΐ ₂ θ ₄ , such as 97% - 99, 9% Zn in the binder is present as ZnAl ₂ 0 ₄ . In preferred embodiments the catalyst has been fully spinelized meaning that all or substantially all of the Zn in the binder is present as ΖηΑΐ ₂ θ ₄ .

ZnO in the binder is active in cracking methanol which is an undesired reaction in MTA. Depending on the means of production and after-treatment of the catalyst more or less of the Zn in the alumina binder may be present as ΖηΑΐ ₂ θ ₄ . Steaming or calcination of a Zn impregnated catalyst as commonly applied in production of metal/zeolite systems may result in a partial spinelization of the Zn (ZnO + AI ₂ O ₃ -> ΖηΑΐ ₂ θ ₄ ) . However, it has been shown that with a high Zn content even a relatively high degree of spinelization may lead to substantial MeOH cracking, but that a very desira ^¬ ble catalyst is achieved with a high degree of or prefera ^¬ bly full spinelization of Zn in the alumina binder i.e. where all or substantially all of Zn in the binder is pre ^¬ sent as ΖηΑΐ ₂ θ ₄ .

A bifunctional catalyst where all of or substantially all Zn in the binder is present as ΖηΑΐ ₂ θ ₄ and where substan ^¬ tially no ZnO is present in the binder as described herein exhibits a low selectivity to CO _x even if the Zn content is high e.g. above 9 wt%. Thus, in preferred embodiments the fresh (start of run) catalyst has a CO _x selectivity (deter ^¬ mined at 420°C, 20 bar, 10 mol% methanol and a WHSV of 1.6) below 8 % preferably below 7 % such as 6% or below, or 5% or lower, or even 2% or lower. The CO _x selectivity is defined as the molar percentage of methanol in the feed converted into CO and CO ₂ according to the net reactions: CH ₃ OH -> CO + 2 H ₂ CH3OH + H ₂ 0 -> C0 ₂ + 3H ₂ Thus, by the present application is provided a preferred bifunctional catalyst comprising alumina binder, H- ZSM-5 and 8 - 15 wt% Zn in the total catalyst and where the Zn in the binder is fully or substantially fully spinelized. Said catalyst provides a high aromatics yield in a MTA reaction while cracking of the methanol is reduced to below 7%.

An exemplary bifunctional catalyst may desirably comprise 30-65 wt% H-ZSM-5, 5-40 wt% ZnAl ₂ 0 ₄ , 0-40 wt% A1 ₂ 0 ₃ , 0-10 wt% ZnO.

The catalyst may further in some embodiments be character ^¬ ized by having 0.1-12 wt% such as 1 - 7 wt% Zn present in the zeolite phase. Alternatively, it may comprise 50-60 wt% H-ZSM-5, 10-35 wt% ZnAl ₂ 0 ₄ , 2-25 wt% A1 ₂ 0 ₃ , 0-7 wt% ZnO. In order to avoid the presence of free ZnO in the binder phase, it may be benefi ^¬ cial to have at least a small excess of A1 ₂ 0 ₃ which is not spinelized in reaction with ZnO. Using a higher amount of A1 ₂ 0 ₃ in the preparation of the "base catalyst" will lead to a more robust catalyst preparation process.

Due to gradual coking of the catalyst during operation the catalyst must be regenerated at intervals in a stream com- prising 0 ₂ . A partially spinelized catalyst with a moderate to high ZnA^C^iZnO content may e.g. be obtained by heating the Zn- impregnated base catalyst at 300-500°C in air. A partially spinelized catalyst with a very high

ΖηΑΐ ₂ θ ₄ ΐ ΖηΟ content, fully spinelized catalyst or a sub ^¬ stantially fully spinelized catalyst may be obtained by heating the Zn impregnated catalyst at 300 - 550°C in steam or in an atmosphere comprising at least 10 vol%, 30 vol% 50 vol% or 80 vol% steam.

A partially spinelized catalyst with a very high

ΖηΑΐ ₂ θ ₄ ΐ ΖηΟ content, fully spinelized catalyst or a sub ^¬ stantially fully spinelized catalyst may be obtained by heating a partially spinelized catalyst at 300 - 550°C in steam or in an atmosphere comprising at least 10 vol%, 30 vol% 50 vol% or 80 vol% steam.

An at least partially spinelized catalyst, preferably a partially spinelized catalyst with a very high ΖηΑΐ ₂ θ ₄ con ^¬ tent, a fully spinelized catalyst or a substantially fully spinelized catalyst as described herein may be provided in numerous ways including obtaining a desired spinelized cat ^¬ alyst during production or by producing a catalyst with a spinelization degree below the desired spinelization percentage and followed by steaming said catalyst in a subse ^¬ quent step e.g. as in an in situ steaming step to obtain a catalyst with a desired degree of spinelization. Various methods may be applied to produce the bifunctional catalyst: The two components (Zn and Zeolite) may consti ^¬ tute an integrated entity, e.g. as obtained by introducing the Zn component by impregnation or ion-exchange to the ze ^¬ olite, either onto the zeolite itself or onto an extrudate in which the zeolite is embedded in an alumina binder. The Zn component may also be added in the form of a salt, ei- ther as a solid or in solution, or an oxide, hydroxide or carbonate together with the zeolite, binder and/or lubricants prior to shaping, e.g. during extrusion or pelletiza- tion . The post-impregnation treatment (calcination or similar heat treatment) is preferably carried out in a humid atmos ^¬ phere, e.g., by heating the Zn impregnated base catalyst at 300 - 550°C in steam or in an atmosphere comprising at least 10 vol%, 30 vol% 50 vol% or 80 vol% steam.

Also physical mixtures of several acidic and metal compo ^¬ nents may be applied and the mixture may be charged to the reactor to form a uniform mixture or to form alternating layers or they may be graded to various degrees.

Methanol to aromatics

The obtained hydrocarbon stream may be rich in aromatics, in which case the process hereinafter is referred to as a methanol-to-aromatics (MTA) process.

In the MTA conversion process methanol and/or other oxygenates is converted into aromatics while dehydrogenation of hydrocarbons including one or more types of hydrocarbons, comprising naphthenes, paraffins and isoparaffins , into olefins and/or aromatics is taking place. The MTA conversion process may preferably be carried out at a pressure 5-60 bar, preferably 10-40 bar, temperature 300- 500°C, preferably 330-480°C and/or weight hourly space ve ^¬ locities (kg alcohol and/or ether feed per kg of catalyst per hour) between 0.1 and 10 preferably 0.3-3.

The MTA process may provide a product particularly suited as feedstock for downstream aromatics processing, e.g. for making para-xylene.

Methanol to gasoline

The obtained hydrocarbon stream from the present process may comprise raw gasoline, in which case the process here- inafter is referred to as a methanol-to-gasoline (MTG) pro ^¬ cess.

In the MTG conversion process methanol is converted into a raw gasoline comprising aromatics, naphthenes, paraffins and isoparaffins and olefins while dehydrogenation of hy ^¬ drocarbons including one or more types of hydrocarbons, comprising naphthenes, paraffins and isoparaffins, into olefins and/or aromatics is taking place. The MTG conversion process may preferably be carried out at a pressure 5-60 bar, preferably 10-40 bar, temperature 300- 500°C, preferably 300-430°C and weight hourly space veloci ^¬ ties (kg alcohol and/or ether feed per kg of catalyst per hour) between 0.1 and 10 preferably 0.5-3.

Hydrogen rejection The dehydrogenation reaction, promoting the formation of aromatics, is equilibrium-limited and in a synthesis loop, hydrogen may build up until the dehydrogenation reaction becomes essentially extinct.

Therefore, removal of hydrogen from the loop ensures that the dehydrogenation reaction is sustained. The hydrogen removal may be continuous, carried out in intervals and/or determined by preselected or adapted desired ¾ concentra- tions e.g. in depleted or partially depleted recycle stream ( s ) .

According to the invention ¾ may be at least partly re ^¬ moved by one or more methods removing ¾ at one or more stages, from one or more streams.

As mentioned above, removal of ¾ from the synthesis loop may be passive, i.e. by purging part of the gaseous recycle stream, or active, to be understood as methods by which one or more recycle streams are selectively depleted or par ^¬ tially depleted with respect to H ₂ .

Depending on the method used for ¾ removal more or less ¾ may be removed. Also, the selectivity in the removal to- wards ¾ may vary. For example, if a permselective membrane is used for the removal the membrane may allow other sub ^¬ stances than H2 such as methane and carbon monoxide to pass through in which case not only ¾ is removed. Similarly, e.g. oxidation steps may result in oxidation of other sub- stances than ¾, e.g. carbon monoxide. According to some embodiments an at least partially ¾ de ^¬ pleted recycle stream is obtained from the gaseous stream by passing the gaseous stream to a hydrogen perm-selective membrane. I.e. the gaseous stream is provided to a stage wherein the gas is passed to a membrane removing ¾ . The permeation of ¾ in the selective membrane may be adjusted to leave a pre-determined amount of hydrogen in the reten- tate recycled to the conversion step as the at least par ^¬ tially H ₂ depleted recycle stream.

Alternatively, the at least partially ¾ depleted recycle stream is obtained from the gaseous stream by passing said gaseous phase, after admixture with a predetermined amount of dioxygen, to a catalytic preferential oxidation, com- monly known as PrOx, step where hydrogen is reacted with said predetermined amount of oxygen to form water and recy ^¬ cling said reacted stream, at least partially depleted in hydrogen, to the conversion step. The catalytic oxidation is preferably carried out at essentially the same pressure as in the conversion step at a temperature between 50 and

300°C, preferably, 100-200°C. Catalysts effective in selec ^¬ tive oxidation typically comprise a noble metal supported on an inorganic metal oxide. Optionally, the at least partially ¾ depleted recycle stream is obtained from the gaseous stream by passing said gaseous phase, after admixture with a predetermined amount of a non-dioxygen hydrogen scavenger, to a catalytic oxidation step where hydrogen is reacted with said predetermined amount oxidizing agent and recycling said reacted stream, at least partly depleted in hydrogen, to the conversion step . The hydrogen scavenger may be an aldehyde or a ketone, in which case said aldehyde and hydrogen is converted into an alcohol over a hydrogenation catalyst. For example, the al- dehyde may be formaldehyde in which case formaldehyde and hydrogen is converted into methanol over a hydrogenation catalyst .

The hydrogen scavenger may be hydrogen peroxide in which case hydrogen peroxide and hydrogen is converted into water over a hydrogenation catalyst.

The separation step, where the effluent from the conversion step is separated into a gas phase and aqueous and hydro- carbon liquid phases, is typically conducted by cooling the conversion effluent essentially at the synthesis pressure, at 10-50 bar, typically by cooling to a temperature between 30 and 80°C. Under these conditions ¾ is almost only found in the gaseous phase in the separator, and the liquid hy- drocarbon (as well as the aqueous phase) is almost com ^¬ pletely ¾ free. This means that the liquid hydrocarbon phase may be separated into a product phase and one or more lower- and/or higher-boiling phases, and at least one of the lower- and/or higher-boiling phases can then at least partially be recycled to the conversion step as one of one or more at least partially ¾ depleted recycle.

Thus, several embodiments according to the present inven ^¬ tion have been proposed, wherein hydrogen may be removed from the synthesis loop and thereby enhance the selectivity to aromatics. Other means of promoting the selectivity to aromatics for example includes increasing the temperature in the conversion reactor, thereby thermodynamically fa ^¬ vouring the dehydrogenation reaction and, in turn, enhancing the selectivity to aromatics. The degree of ¾ removal affects the products from the con ^¬ version. As mentioned above, if no means of hydrogen re ^¬ moval is established, hydrogen will accumulate in the loop and, eventually, the dehydrogenation reaction will cease. Therefore, continuous removal of hydrogen from the loop may be required in order to sustain the dehydrogenation reaction. To which extent yield and selectivity may be in ^¬ creased depends on how much hydrogen is rejected from the synthesis loop. Thus, if only a limited part of the hydro ^¬ gen is rejected, e.g. purging part of the recycle gas (pas- sive rejection) this will lead to a product with a slight enrichment in aromatics, whereas a high degree of hydrogen rejection, as may be achieved by active ¾ rejection as de ^¬ scribed above, will provide a product with a high aromatics content and even with aromatics as the main constituent.

The extent of hydrogen removal, therefore, represents a convenient and efficient tool for controlling aromatics se ^¬ lectivity and product yield. In a methanol-to-gasoline context removal of hydrogen, pas ^¬ sive or active, will lead to a higher gasoline yield as well as improved octane due the enhanced selectivity to ar ^¬ omatics. However, in most areas the aromatics content is regulated, typically to 35 vol% (some 40 wt%) , which im- poses a limit to how much hydrogen may be removed. Nonethe ^¬ less, a few per cent improvement in gasoline yield and at the same time gaining a few numbers in octane represents a significant economic advantage. Moreover, as will be shown by examples, the selectivity to heavy oil (carbon number Cll and higher) and durene (1,2,4,5 tetramethylbenzene) is significantly reduced by the present process. I.e. the in- vention provides for making gasoline with improved yield and octane and with reduced selectivity to heavy oil and durene .

In an MTA context, there is no limit as to how much hydro- gen should be removed and aromatic yields in the order of 70% or more may be achieved at close to complete hydrogen rej ection .

The selectivity to aromatics increases with temperature, in part due to higher rates of reaction, but primarily, because dehydrogenation becomes favoured thermodynamically by an increase in temperature. Consequently, higher aromatic yields may be obtained by conducting the conversion at higher temperatures.

Thus the present process provides a method for making hy ^¬ drocarbons wherein the content of aromatics in the product slate may be controlled to a significant extent by applying a catalyst or a combination of catalysts active in dehydro- genation and methanol-to-hydrocarbons formation and a means of passive or active rejection of hydrogen from the synthe ^¬ sis loop, the degree of which passive and/or active, con ^¬ trols the amount of aromatics in the product slate. The feed stream may comprise methanol, dimethyl ether or mixtures hereof. In preferred setups the process is carried out in a fixed bed reactor.

Thus, in summary a highly advantageous process is achieved by the present invention wherein the advantages of the ¾ removal from the process loop is fully utilized together with the bifunctional catalyst preferably having a high Zn content e.g. >8 wt% as well as a high degree of spineliza- tion of the Zn in the alumina binder and thereby a reduced COx selectivity resulting in a reduced ¾ formation from

MeOH cracking. I.e. in some aspects of the present process is provided a process for production of hydrocarbons with increased aromatics content due to the ¾ reducing features of the process and catalyst.

Example 1 : Preparation of catalyst

A base catalyst containing 65 wt % H-ZSM-5 and 35% AI2O3 was prepared by mixing followed by extrusion following well known pro- cedures . Upon calcination, samples of the base catalyst were impregnated with an aqueous solution containing zinc nitrate at different Zn concentrations. The resulting pore-filled extru- dates were heated to 470°C in air and kept at 470°C for 1 h to obtain catalysts with various amounts of Zn.

Example 2 : Catalyst activity and regeneration

Catalysts prepared by the procedure described in example 1 were subjected to conversion of methanol at 420°C in an isothermal fixed bed reactor. 2 was used as an inert co-feed to obtain a methanol concentration of 7 mol% in the reactor inlet. The total pressure was 20 bar, and the space velocity (WHSV) of methanol was 2 h ^_1 . Zn/H-ZSM-5 catalysts suffer from reversible as well as irreversible deactivation. Deposition of carbon (coke) on the catalyst is responsible for reversible deactivation. In the example shown in table 1, the deactivated (coked) catalyst is regenerated by removal of the deposited carbon by combustion in a flow of 2% 0 ₂ (in N ₂ ) at 500°C.

Due to irreversible deactivation, the catalyst did not fully regain its activity after regeneration. The results in table 1 show, that a catalyst containing 10% Zn is able to regain significantly more of its original activity after regeneration than a catalyst containing 5% Zn.

Table 1: Catalyst activity after regeneration. Wt% of aromatics in hydrocarbon product is defined as the mass of aromatics relative to the total mass of hydrocarbons in the effluent stream.

Example 3 : Stability towards steaming

To simulate catalyst activity after extended operation under industrial conditions, the catalysts were subjected to methanol conversion after steaming under severe conditions. Methanol conversion was performed under the same conditions as in example 2. The results in Table 2 show that the catalyst containing 10% Zn retains significantly more of its original activity than the catalyst containing 5 wt% Zn after severe steaming. Table 2: Loss of catalyst activity upon severe steaming (100% steam for 48h at 500°C and 1 bar) . Wt% of aromatics in hydrocarbon product is defined as the mass of aromatics relative to the total mass of hydrocarbons in effluent stream.

Example 4 : Methanol cracking vs . Zn content

Cracking (decomposition) of methanol/DME can occur via several mechanisms. For example, the acidic sites in the catalyst may catalyze cracking of DME to CH4, CO, and H2, while certain Zn species catalyze cracking of methanol to CO and H2. CO2 can be formed as a primary cracking product or indirectly via the water gas shift reaction.

When methanol is converted over a catalyst containing Zn, part of the methanol is converted to COx due to cracking, which results in lower yield of hydrocarbon products . Methanol conversion has been performed at 420°C, 20 bar, 10mol% methanol (N2 balance), and a space velocity (WHSV) of 1.6.

The results in Table 3 were obtained using catalysts prepared according to example 1. The results show that the cracking activity is highly dependent on the amount of Zn, i.e. higher Zn content leads to higher cracking activity. Table 3: CO _x selectivity at different contents of Zn

Example 5: COx selectivity after calcination and steaming

A base catalyst containing 65% ZSM-5 and 35% AI2O3 was impregnated with aqueous zinc nitrate solution. The resulting pore filled extrudates were calcined in air and steam, respectively. Furthermore, the catalyst calcined in air was subjected to steaming after calcination. Methanol conversion over these cata lysts was performed using the same conditions as in example 4.

The results in table 4 show that the presence of steam during calcination of the impregnated catalyst or heating the catalyst in the presence of steam after calcination leads to lower selec tivity to COx. This observation may be rationalized by the fact that the presence of steam leads to formation of ΖηΑΐ2θ4 rather than free ZnO in the binder phase.

Table 4: CO _x selectivity for catalysts containing 10% Zn, calcined in the presence of different amounts of steam

Condition CO _x selectivity

(%)

Calcined in air 9

Calcined in steam (500°C, 2h) 2

Calcined in air, steamed after calcination 4

(500°C, 5 h) Calcined in air, steamed after calcination <0.1

(500°C, 48 h)

Examples 6-8: Examples 6-8 demonstrate the influence of hydrogen on yield and selectivity. A reference experiment with un-doped H- ZSM-5 (example 6) was conducted at 420°C and a weight hourly space velocity of 1.6. Example 6 was then repeated, except that a Zn-impregnated catalyst containing 2.9 wt% zinc was applied (example 7) . Finally, example 7 was re ^¬ peated, but this time the nitrogen carrier gas was replaced by a mixture of 5 vol% hydrogen in nitrogen (example 8) .

The results of the three experiments, averaging a period of 30 hours are reported in Table 5.

Table 5

Example 6 7 8

Temperature (°C) 420 420 420

Pressure (bar g) 20 20 20

WHSV 1.6 1.6 1.6

Carrier gas (vol% H ₂ / ₂ ) 0/100 0/100 5/95

Average over period (hrs) 0-30 0-30 0-30

Products (wt% of total HC)

n-paraffins 27.30 12.45 14.31 i-paraffins 27.86 24.52 29.89

Olefins 7.92 13.79 9.04

Naphthenes 1.30 2.56 2.94

Aromatics 35.18 46.50 43.72 C11+ ("heavy oil") 0.61 0.58 0.38

Overall molar H/C ratio in

hydrocarbon product 2.01 1.83 1.89

50.50 65.98

Overall C5+ yield 66.99

Aromatics in C5+ fraction 69.91 69.38 66.25

Olefins in C5+ fraction 1.76 3.76 2.11

Durene in C5+ fraction 1.48 2.54 1.96

By comparing example 7 with reference example 6 it is again shown that incorporation of Zn to the H-ZSM-5 catalyst leads to a significant improvement in yield and aromatics selectivity. The effect of hydrogen at the inlet of the re ^¬ actor is demonstrated by comparing example 7 and 8, showing that the presence of hydrogen causes a decline in the aro ^¬ matics selectivity and a slight reduction in the C5+ yield. However, when comparing example 8 with reference example 6 it is clear that, despite the presence of hydrogen, the C ₅₊ yield and aromatics selectivity is still significantly higher, showing that the Zn-doped catalyst is capable of improving yield and aromatics selectivity even when hydro ^¬ gen is present at the reactor inlet. This demonstrates that, in a synthesis loop, controlling the amount of hydro ^¬ gen recycled to the conversion reactor inlet provides a means of directing the selectivity towards aromatics.

It is also observed (example 7 & 8) that the presence of hydrogen reduces the formation of durene and heavy oil. At the high temperatures applied in examples 6-8 some increase in durene levels are observed when applying the Zn-doped catalysts relative to the un-doped catalyst. However, in these examples durene selectivity are already so low (below 4 wt%) that it is without any significance to driveability.

In conclusion, examples 6 - 8 show that the incorporation of Zn to the zeolite increases the aromatics selectivity and the C5+ yield, even so in the presence of hydrogen in the feed and that the presence of hydrogen in the feed leads to a reduction in the formation of durene and heavy oil. Thus, in a synthesis loop, finite amounts of hydrogen in the recycle stream is not necessarily prohibitive in re ^¬ gard to achieving the desired effects according to the pre ^¬ sent invention, namely that of increasing C5+ yield and ar ^¬ omatics . In the following the process and plant is further described with reference to the accompanying drawings. The drawings show exemplary embodiments of the present process and plant and are not to be construed as limiting to the scope of the present application.

Fig. 1 shows a schematic setup of a first embodiment ac ^¬ cording to the present invention,

Fig. 2 shows a schematic setup of a second embodiment ac- cording to the present invention, and

Fig. 3 shows a schematic setup of third embodiment accord ^¬ ing to the present invention. Fig. 1 shows a schematic overview 1 of a plant/process ac ^¬ cording to the present application. A conversion effluent 2 is fed to a first separator 3 wherein the conversion effluent is separated into three streams: process condensate 4, first product stream 5 and a gas stream 6. The first product stream is fed to a second separator 7 wherein the first product stream is separated into a second product stream 8 and a LPG stream comprising C3 - C4 where- from the recycle stream 9a is taken. As described herein the first product stream is near ¾ free and ¾ is present predominantly in the gas phase in the first separator. Thus when the first product stream is separated into the second product stream 8 and the recycle stream the recycle stream obtained is a ¾ depleted recy- cle.

The second product stream 8 may be send to further pro ^¬ cessing, upgrade, storage etc. Fig. 2 shows a schematic overview 10 of a plant/process ac ^¬ cording to the present application. A conversion effluent 2 is fed to a first separator 3 wherein the conversion effluent is separated into three streams: process condensate 4, first product stream 5 and a gas stream 6. At least part of the gas stream 6 is taken through a ¾ depletion step 11 whereby a ¾ depleted recycle stream 9b obtained.

The ¾ depletion step 11 may comprise e.g. a H2 permeable membrane and/or a catalytic oxidation step.

Fig. 3 shows a schematic overview 12 of a plant/process wherein the embodiments of fig. 1 and 2 are combined whereby a ¾ depleted stream is obtained from the gas stream 9b and from the LPG fraction taken from the second separator as H ₂ depleted stream 9a. In the embodiments of fig. 1, 2 and 3 the one or more ¾ depleted streams are returned to the conversion step (not shown) from which the conversion effluent 2 is obtained. The conversion step may be a MTG or a MTA process.