# Patent

## Title
Process for producing olefins

## Abstract

The present invention provides a process for producing olefins from a feed comprising at least methane, ethane and carbon dioxide. The feed is separated into at least a methane-comprising feed, an ethane-comprising feed and a carbon dioxide-comprising feed. At least part of the methane-comprising feed is converted to a synthesis gas. The ethane-comprising feed is cracked to obtain at least olefins and hydrogen. At least part of the carbon dioxide-comprising feed and at least part of the synthesis gas are used to synthesis oxygenates. At least part of the oxygenates are converted in an oxygenate-to-olefin (OTO) zone to obtain at least olefins and hydrogen. At least part of the cracking effluent and at least part of the OTO zone effluent are combined to obtain a combined effluent from which hydrogen is separated. At least part of the hydrogen is supplied to the oxygenate synthesis zone.

## Background
This application claims the benefit of European Application No. 09175607.2 filed Nov. 10, 2009 which is incorporated herein by reference.
FIELD OF THE INVENTION
The invention relates to a process for producing olefins and an integrated system for producing olefins.
BACKGROUND OF THE INVENTION
In recent years increasing attention is given to the exploration and utilisation of natural gas resources around the globe. A disadvantage of natural gas with respect to oil is the difficulty to transport large volumes of natural gas from the source to the market. One way of efficiently transporting natural gas is by liquefying the natural gas and to transport the liquefied natural gas (LNG). Another way is to convert the methane in the natural gas to liquid hydrocarbons using a Gas-to-Liquid process (GtL). The GtL products are typically liquid and can be transported in a similar way as traditional oil and oil products.
Besides methane, the natural gas typically comprises other hydrocarbons such as ethane, propane, and butanes. Such a natural gas is referred to as wet gas. The latter two can be added to the LPG pool, however, ethane cannot. Moreover, for various reasons the ethane content in the natural gas supplied to an LNG or GtL process is restricted and therefore a significant part of the ethane must be removed from the natural gas prior to providing the natural gas to either a LNG or GtL process.
Although, the application of ethane is limited, typically ethane is combusted in a furnace to provide heat; its corresponding olefin ethylene is a base chemical with a wide application and is of great commercial interest. Ethane can be converted into ethylene, e.g. using a thermal cracking process. Subsequently, the ethylene can be used to produce e.g. polyethylene, styrene or mono-ethylene-glycol. The conversion of ethane to ethylene is highly endothermic and requires significant energy input. In addition, the capex for the ethane to ethylene process, in particular the back-end work-up section, and the subsequent ethylene conversion processes is high and a minimum ethylene production capacity is required to make it economically benign.
When, the ethane content in the natural gas is too low, and consequently insufficient ethane is available, the ethane/ethylene route becomes unattractive.
This problem becomes even more pronounced, in the case where the natural gas is withdrawn from relatively small reservoirs, especially those located in remote isolate locations, also revered to as stranded natural gas. Of course, this stranded natural gas may be converted to LNG or GtL products. However, this requires the stranded gas reservoir to sustain a minimum production level per day in order to make the investments worthwhile. Typically, such stranded natural gas reservoirs cannot achieve sufficient production levels to sustain a GtL or LNG plant. In addition, insufficient ethane is co-produced to sustain an ethane to ethylene process and subsequent ethylene conversion processes.
It has been suggested to combine an ethane steam cracker with Oxygenate-to-Olefin (OTO) processes, which can produce additional ethylene. For instance, C. Eng et al. (C. Eng, E. Arnold, E Vora, T. Fuglerud, S. Kvisle, H. Nilsen, Integration of the UOP/HYDRO MTO Process into Ethylene plants, 10th Ethylene Producers' Conference, New Orleans, USA, 1998) have suggested to combine UOP's Methanol-to-Olefins (MTO) process with a naphtha or ethane fed steam cracker. It is mentioned that by combining both processes sufficient ethylene can be produced, while coproducing valuable propylene. A disadvantage mentioned by C. Eng et al. is the fluctuating price of methanol, which is the primary feed to the MTO reaction.
In WO 2009/039948 A2, a combined stream cracking and OTO process is suggested for preparing ethylene and propylene. According to WO 2009/039948 A2, in this process, a particular advantage is obtained by combing the back-end of both processes. The methanol feedstock is produced from methane, requiring a sufficient supply of methane.
In US2005/0038304, an integrated system for producing ethylene and propylene from an OTO system and a steam cracking system is disclosed. According to US2005/0038304, in this process, a particular advantage is obtained by combining the back-end of both processes. The methanol feedstock to the OTO process is produced from synthesis gas. However, according to US2005/0038304 the production of methanol from synthesis gas has high energy requirements due to the endothermic nature of the synthesis gas production process, such an endothermic synthesis gas production process is normally steam methane reforming.
In US2002/0143220A1, a process is described for producing olefins. A hydrocarbon feedstock is oxidatively dehydrogenated to produced olefins and synthesis gas. The synthesis gas may be converted to methanol. The methanol may be converted to ethylene.
A problem with using natural gas to produce ethylene is that together with the hydrocarbons in the natural gas a substantial amount of carbon dioxide may be co-produced from the subsurface natural gas or oil reservoir. Such carbon dioxide is also referred to as field carbon dioxide. Some subsurface natural gas or oil reservoirs comprise substantial concentrations of carbon dioxide, up to 70 mol % based on the total content of the gas extracted from the reservoir. This carbon dioxide must be sequestrated or otherwise captured and stored.
An additional problem is that, especially when the carbon dioxide content is high, not enough ethane is produced from the subsurface reservoir to sustain a ethane cracker with full work-up section.
In US2007/049647A1 it is described that carbon dioxide co-produced with natural gas may be mixed with synthesis gas to produce a methanol feed for a methanol-to-olefins process.
In WO2007/142739A2, a process is described for producing methanol from synthesis gas. The methanol may be used for producing olefins. In the process described in WO2007/142739A2, a hydrogen stream comprising greater than 5 mol % methane is combined with the synthesis gas. The hydrogen stream may for instance be obtained from a steam cracking process. According to WO2007/142739A2 it is desirable that the amount of CO2 in the synthesis gas feedstock is minimized to reduce the need for later separation of water from the crude methanol.
There is a need in the art for a process for producing ethylene from a feed comprising, besides hydrocarbons, carbon dioxide, wherein the amount of carbon dioxide that needs to be sequestrated or otherwise captured and stored is further reduced.
SUMMARY OF THE INVENTION
It has now been found that it is possible to further reduce the amount of carbon dioxide that needs to be sequestrated or otherwise captured and stored, when producing ethylene from a feed comprising, besides hydrocarbons, carbon dioxide by using an integrated process, wherein at least part of the field carbon dioxide together with hydrogen produced internally by the process is used to synthesis oxygenates that can be used as feedstock to an OTO process.
Accordingly, the present invention provides a process for producing olefins, comprising:
a. providing a feed comprising at least methane, ethane and carbon dioxide;
b. separating the feed into at least a methane-comprising feed, an ethane-comprising feed and a carbon dioxide-comprising feed;
c. providing at least part of the methane-comprising feed to a process for preparing synthesis gas to obtain a synthesis gas;
d. cracking the ethane-comprising feed in a cracking zone under cracking conditions to obtain a cracking zone effluent comprising at least olefins and hydrogen;
e. providing at least part of the carbon dioxide-comprising feed and at least part the synthesis gas obtained in step c) to an oxygenate synthesis zone and synthesising oxygenates;
f. converting at least part of the oxygenates obtained in step (e) in an oxygenate-to-olefin zone to obtain an oxygenate-to-olefin effluent comprising at least olefins and hydrogen;
g. combining at least part of the cracking zone effluent and at least part of the OTO zone effluent to obtain a combined effluent;
h. separating hydrogen from the combined effluent and providing at least part of the hydrogen to the oxygenate synthesis zone in step (e).



BRIEF DESCRIPTION OF THE DRAWING
FIG. 1 is a schematic representation of an integrated system for producing olefins according to the invention.



DETAILED DESCRIPTION OF THE INVENTION
The process according to the present invention is directed at producing olefins, in particular lower olefins (C2-C4), more in particular ethylene and propylene. Several processes, such as ethane cracking or Oxygenate-to-Olefin (OTO) processes can produce olefins. The ethane cracking and the OTO processes typically produce ethylene from different starting materials. In the case of the ethane cracking step, the feed is preferably an ethane-comprising feed. The OTO step on the other hand uses an oxygenate-comprising feed. Preferred oxygenates include alkylalcohols and alkylethers, more preferably methanol, ethanol, propanol and/or dimethylether (DME), even more preferably methanol and/or dimethylether (DME).
By using an integrated process comprising both an ethane cracking process and an OTO process it is possible to produce lower olefins from a mixed feed comprising both methane and ethane. Such a feedstock is first separated into a methane-comprising feed and an ethane-comprising feed. The ethane-comprising feed is used as feed to the ethane cracking process. The methane-comprising feed is converted to synthesis gas, which is used to synthesize the oxygenate-comprising feed required for the OTO process.
The most common feedstock comprising both methane and ethane is natural gas. Besides methane and ethane, the natural gas may contain significant amounts of carbon dioxide. This carbon dioxide is contained in the subsurface reservoir from which the natural gas is produced. As the natural gas is withdrawn from the reservoir the carbon dioxide is co-produced as part of the natural gas.
There is a desire not to emit this coproduced carbon dioxide to the atmosphere, therefore it must be sequestrated or otherwise captured and stored.
It has now been found that by producing olefins by an integrated process comprising an ethane cracking step and an Oxygenate-to-Olefin (OTO) step, a synergetic use of the feedstock comprising at least methane, ethane and carbon dioxide is achieved by using at least part of the carbon dioxide that is in the feedstock, further also referred to as field carbon dioxide for synthesising oxygenates, preferably by converting hydrogen with carbon monoxide and/or carbon dioxide into methanol and/or dimethylether. Wherein at least part of the hydrogen is co-produced in the ethane cracking step and OTO step.
These oxygenates are subsequently fed to the OTO reaction to form further olefins.
By using field carbon dioxide the need to sequestrate or otherwise capture and store the field carbon dioxide is reduced if not omitted. At least part of the field carbon dioxide is used to produce at least part of the oxygenate feed to the OTO zone. As a result, the field carbon dioxide is no longer dissipated into the atmosphere or otherwise captured and stored, but rather used to produce valuable oxygenates. Thereby, the carbon dioxide penalty for producing the olefins is reduced.
In addition, the field carbon dioxide preferably does not comprise significant amounts of inerts such as N2 or Ar. These inerts may typically be present in the natural gas or purified oxygen provided to produce synthesis gas for methanol production. By providing field carbon dioxide as part of the feed to the oxygenate synthesis zone the inert content in this feed may be reduced.
By providing at least part of the field carbon dioxide to the oxygenate synthesis zone the amount of synthesis gas needed to synthesise the oxygenates may be reduced. The synthesis gas is typically produced by partial oxidation of hydrocarbons using essentially pure oxygen or at least oxygen enriched air. The production of pure oxygen is highly energy consuming, therefore a reduction in the synthesis gas demand also reduces the oxygen demand, which in return results in decreased energy consumption and carbon dioxide production. Moreover, the capital investment can be reduced as a smaller oxygen production unit is required.
In the process according to the invention a feedstock comprising at least methane, ethane and carbon dioxide is provided. In step (b) of the process, the feedstock is separated into at least a methane-comprising feed, an ethane-comprising feed and a carbon dioxide-comprising feed. Hereinbelow the carbon dioxide-comprising feed will also be referred to as carbon dioxide ex. field.
In step (c) of the process, at least part of the methane-comprising feed is provided to a process for preparing synthesis gas. Such processes for preparing synthesis gas preferably include non-catalytic partial oxidation processes, catalytic partial oxidation processes, steam methane reforming processes, auto-thermal reforming processes, and water-gas-shift processes, wherein hydrocarbons, in particular methane are converted with oxygen and optionally steam to at least hydrogen and preferably carbon monoxide. Although, a water-gas-shift process is in principle not a process for preparing a synthesis gas, the effluent of a water-gas-shift process typically comprises carbon monoxide, carbon dioxide and hydrogen. The methane-comprising feed may be provided to several processes for preparing synthesis gas.
The methane-comprising feed may be converted into a synthesis gas having a hydrogen and carbon monoxide and/or carbon dioxide molar ratio, as defined hereinbelow:

molar ratio=(#mol H2−#mol CO2)/(#mol CO+#mol CO2),

Preferably, the methane-comprising feed is converted in to a synthesis gas having a hydrogen and carbon monoxide and/or carbon dioxide molar ratio, which is above the ratio preferred for synthesising methanol, i.e. sources that are carbon deficient.
As can be seen in the definition of the molar ratio hereinabove, a water-gas-shift does not influence the molar ratio obtained.
Preferably, a synthesis gas is used comprising little or no carbon dioxide. When using a synthesis gas with little or no carbon dioxide, more carbon dioxide ex. field can be added. Low carbon dioxide comprising synthesis gases are therefore preferred, more preferred are synthesis gases comprising carbon dioxide in the range of from 0 to 8 mol % even more preferably of from 0.1 to 3 mol %, still even more preferably 0.2 to 2.5 mol %, based on the total number of moles in the synthesis gas.
Such low carbon dioxide synthesis gases are preferably produced by non catalytic partial oxidation processes for preparing synthesis gas. A partial oxidation catalyst typically induces some water-gas-shift in the presence of water. As a result carbon monoxide is shifted to carbon dioxide. An additional advantage is that non-catalytic partial oxidation processes do not require the addition of substantial amounts of water to the process. Such as non-catalytic partial oxidation processes. Processes producing substantial amounts of carbon dioxide include for instance Steam Methane Reforming. Therefore, the use of a synthesis gas from a Steam Methane Reforming process is less preferred.
In step (d) of the process according to the invention the ethane-comprising feed is provided to a cracking zone and cracked under cracking conditions. The obtained cracking product comprises olefins and hydrogen, in particular at least ethylene and hydrogen.
In step (e) of the process according to the invention at least part of the carbon dioxide ex. field and at least part the synthesis gas obtained in step c) are provided to an oxygenate synthesis zone. Next to carbon dioxide ex. field and synthesis gas obtained in step c) other suitable sources of hydrogen, carbon monoxide and or/carbon dioxide may be provided to the oxygenate synthesis zone of step (e).
In the oxygenate synthesis zone, oxygenates are preferably synthesised by reacting hydrogen with carbon dioxide and optionally carbon monoxide. Preferably, hydrogen is reacted with carbon dioxide and optionally carbon monoxide to form an oxygenate suitable to serve as a feed to an OTO process, in particular alkylalcohols and alkylethers, preferably methanol and/or di-methyl-ether (DME).
Methanol may be produced directly from hydrogen and carbon dioxide in the oxygenate synthesis zone. Hydrogen can react with carbon dioxide to produce methanol following:

CO2+3H2→CH3OH+H2O.

Preferably, the synthesis gas provided to the oxygenate synthesis zone also comprises carbon monoxide. The hydrogen may than also react with carbon monoxide to also form methanol following:

CO+2H2→CH3OH.

Preferably, the carbon dioxide, hydrogen and preferably carbon monoxide are provided to the oxygenate synthesis zone in a molar ratio in the range of from 2.0 to 3.0, preferably 2.0 to 2.2. The molar ratio herein is defined as:

molar ratio=(#mol H2−#mol CO2)/(#mol CO+#mol CO2).

In case a carbon monoxide is present to convert the hydrogen to methanol, it is preferred that the carbon dioxide concentration in the hydrogen, carbon monoxide and carbon dioxide mixture is in the range of from 0.1 to 25 mol %, preferably 3 to 15 mol %, more preferably of from 4 to 10 mol %, based on the total number of moles hydrogen, carbon monoxide and carbon dioxide in the mixture. The carbon dioxide content, relative to that of CO, in the syngas should be high enough so as to maintain an appropriately high reaction temperature and to minimize the amount of undesirable by-products such as paraffins. At the same time, the relative carbon dioxide to carbon monoxide content should not be too high so that the reaction of carbon dioxide with hydrogen yields less methanol based on the hydrogen provided to the oxygenate synthesis zone. In addition, the reaction of carbon dioxide with hydrogen yields water. If present in too high a concentration, water may deactivate the oxygenate synthesis catalyst.
In the oxygenate synthesis zone the hydrogen, carbon dioxide and preferably carbon monoxide are converted to methanol in the presence of a suitable catalyst. Such catalysts are known in the art and are for instance described in WO 2006/020083, which is incorporated herein by reference. Suitable catalysts for the synthesis of methanol from hydrogen, carbon dioxide and preferably carbon monoxide include:

    
    
        An oxide of at least one element selected from the group consisting of copper, silver, zinc, boron, magnesium, aluminium, vanadium, chromium, manganese, gallium, palladium, osmium and zirconium.
        Preferably, the catalyst is a copper and zinc based catalyst, more preferably in the form of copper, copper oxide, and zinc oxide.
        A copper based catalyst, which includes an oxide of at least one element selected from the group consisting of silver, zinc, boron, magnesium, aluminium, vanadium, chromium, manganese, gallium, palladium, osmium and zirconium.
        Preferably, the catalyst contains copper oxide and an oxide of at least one element selected from the group consisting of zinc, magnesium, aluminium, chromium, and zirconium.
        A catalyst selected from the group consisting of: copper oxides, zinc oxides and aluminium oxides. More preferably, the catalyst contains oxides of copper and zinc.
        A catalyst comprising copper oxide, zinc oxide, and at least one other oxide.
    
    


Preferably, the at least one other oxide is selected from the group consisting of zirconium oxide, chromium oxide, vanadium oxide, magnesium oxide, aluminum oxide, titanium oxide, hafnium oxide, molybdenum oxide, tungsten oxide, and manganese oxide.
Particular suitable catalysts include catalysts comprising in the range of from 10 to 70 wt % copper oxide, based on total weight of the catalyst. Preferably, comprising in the range of from 15 to 68 wt % copper oxide, and more preferably of from 20 to 65 wt % copper oxide, based on total weight of the catalyst.
Such catalyst may preferably also contain in the range of from 3 to 30 wt % zinc oxide, based on total weight of the catalyst. Preferably, it contains in the range of from 4 to 27 wt % zinc oxide, more preferably from 5 to 24 wt % zinc oxide, based on total weight of the catalyst.
Catalyst comprising both copper oxide and zinc oxide, preferably comprise copper oxide and zinc oxide in a ratio of copper oxide to zinc oxide which may vary over a wide range. Preferably, such catalyst comprises copper oxide to zinc oxide in a Cu:Zn atomic ratio in the range of from 0.5:1 to 20:1, preferably from 0.7:1 to 15:1, more preferably from 0.8:1 to 5:1.
The catalyst can be prepared according to conventional processes. Examples of such processes can be found in U.S. Pat. Nos. 6,114,279; 6,054,497; 5,767,039; 5,045,520; 5,254,520; 5,610,202; 4,666,945; 4,455,394; 4,565,803; 5,385,949, with the descriptions of each being fully incorporated herein by reference.
Methanol may be synthesised in the oxygenate synthesis zone by any conventional methanol synthesis process. Examples of such processes include batch processes and continuous processes. Continuous processes are preferred.
Tubular bed processes and fluidized bed processes are particularly preferred types of continuous processes.
The methanol synthesis process is effective over a wide range of temperatures. Preferably, methanol is synthesised in the oxygenate synthesis zone by contacting the hydrogen, carbon dioxide and preferably carbon monoxide with the catalyst at a temperature in the range of from 150 to 450° C., more preferably from 175 to 350° C., even more preferably from 200 to 300° C.
The methanol synthesis process is effective over a wide range of pressures. Preferably, the methanol is synthesised by contacting the hydrogen, carbon dioxide and preferably carbon monoxide with the catalyst in the oxygenate synthesis zone at a pressure in the range of from 15 to 125 atmospheres, more preferably from 20 to 100 atmospheres, more preferably from 25 to 75 atmospheres.
For methanol synthesis, gas hourly space velocities in the oxygenate synthesis zone vary depending upon the type of continuous process that is used. Preferably, gas hourly space velocity of flow of gas through the catalyst bed is in the range of from 50 hr−1 to 50,000 hr−1. Preferably, gas hourly space velocity of flow of gas through the catalyst bed is in the range of from about 250 hr−1 to 25,000 hr−1, more preferably from about 500 hr−1 to 10,000 hr−1.
A methanol synthesis process as described hereinabove may produce several oxygenates as by-products, including aldehydes and other alcohols. Such by-products are also suitable reactants in the OTO reaction. Other less desirable by-products may be removed from the effluent of the oxygenate synthesis zone effluent if required prior to providing the oxygenate synthesis zone effluent to the OTO zone as to form at least part of the oxygenate feed.
Another suitable and preferred oxygenate, which may be synthesised in the oxygenate synthesis zone is dimethylether (DME). DME can be directly synthesized from synthesis gas, but is preferably synthesized from methanol. Optionally, DME is obtained from methanol and hydrogen, carbon dioxide and preferably carbon monoxide. The conversion of methanol to DME is known in the art. This conversion is an equilibrium reaction. In the conversion the alcohol is contacted at elevated temperature with a catalyst. In EP-A 340 576 a list of potential catalysts are described. These catalysts include the chlorides of iron, copper, tin, manganese and aluminium, and the sulphates of copper, chromium and aluminium. Also oxides of titanium, aluminium or barium can be used. Preferred catalysts include aluminium oxides and aluminium silicates. Alumina is particularly preferred as catalyst, especially gamma-alumina. Although the methanol may be in the liquid phase the process is preferably carried out such that the methanol is in the vapour phase. In this context the reaction is suitably carried out at a temperature of 140 to 500° C., preferably 200 to 400° C., and a pressure of 1 to 50 bar, preferably from 8-12 bar, the exact choice depends on the acidity of the catalyst. In view of the exothermic nature of the conversion of methanol to DME the conversion is suitably carried out whilst the reaction mixture comprising the first catalyst is being cooled to maximize DME yield. Suitably, the methanol to DME reaction takes place in a separate section of the oxygenate synthesis zone.
In the case where part of the methanol synthesized is converted into DME, the effluent of the oxygenate zone may comprise methanol and DME in any ratio. Preferably, the ratio of DME to methanol weight ratio is in the range of from 0.5:1 to 100:1, more preferably from 2:1 to 20:1. Suitably the methanol to DME conversion reaction is led to equilibrium. This includes that the DME to methanol weight ratio may vary from 2:1 to 6:1. Evidently, the skilled person may decide to influence the equilibrium by applying different reaction conditions and/or by adding or withdrawing any of the reactants.
In the process according to the invention at least part of the oxygenate feed is methanol and/or DME produced by reacting field carbon dioxide with at least a synthesis gas.
In step (f) of the process according to the invention at least part of the oxygenates obtained in step (e) are provided to an oxygenate-to-olefin zone and converted to obtain at least olefins. Next to the oxygenates obtained in step (e), other oxygenates may be provided to the oxygenate-to-olefin zone of step (f).
In step (d) of the process, olefins, preferably including ethylene, are produced. Typically the olefins, together with any other compounds, leave the cracking zone as cracking zone effluent comprising olefins, including ethylene. Preferably, the ethylene is separated from the remainder of the cracking zone effluent, such as for instance hydrogen and propylene, prior to being used in any subsequent process. The ethylene may be separated using any suitable means in the art. Reference is made to for instance US2005/0038304.
In step (f) of the process also olefins, preferably including ethylene, are produced. The olefins typically leave the OTO zone as part of an OTO zone effluent. As for the ethylene produced in step (d), it is preferred to separate the ethylene for the remainder to the OTO zone effluent prior to providing any of the ethylene to any subsequent process.
In the process according to the invention at least part of the cracking zone effluent and the OTO zone effluent are combined to form a combined effluent prior to separating of the ethylene. It has been found advantageous to combine at least part of the cracking zone effluent and the OTO zone effluent to form a combined effluent. By separating the ethylene from the combined effluent, rather than from the separate effluents, efficient use can be made of the back-end facilities of the processes.
By combining the cracking zone effluent and the OTO zone effluent a further advantage is obtained. In step (d) of the process, hydrogen is produced in significant amounts. However, in step (f) of the process some hydrogen is also produced. The OTO zone effluent therefore also comprises small amounts of hydrogen, typically in the range of from 0.05 to 1 wt % based on the total hydrocarbon content in the OTO zone effluent. The amount of hydrogen in the OTO zone effluent, however, is relatively small, making separating of the hydrogen from the remainder of the OTO zone effluent not worthwhile. By combining at least part of the cracking zone effluent and at least part of the OTO zone effluent to a combined effluent and subsequently separating hydrogen from that combined effluent it is possible to retrieve not only part or all of the hydrogen in the cracking zone effluent, but also at least part of the hydrogen in the OTO zone effluent. The hydrogen obtained from the combined effluent is further also referred to a hydrogen ex. combined effluent. The hydrogen may be separated using any suitable means known in the art, for example cryogenic distillation, pressure swing absorption whereby hydrogen in the hydrogen containing stream absorbs preferentially over the impurities or via hydrogen permeable membrane. By combining the effluents of the cracking zone and the OTO zone, the hydrogen in the OTO zone effluent is also retrieved. Thereby, the amount of hydrogen available to add to the field carbon dioxide and synthesis gas provided to the oxygenate synthesis zone to synthesise oxygenates is increased. Consequently, more field carbon dioxide can be provided to the oxygenate synthesis zone and thus more carbon dioxide is captured in the from of methanol, ethylene and other products.
At least part of the hydrogen obtained from the combined effluent is provided to the oxygenate synthesis zone of step e) and is converted with at least field carbon dioxide and synthesis gas to oxygenates in the oxygenate synthesis zone. This has the advantage that even more field carbon dioxide can be captured. The hydrogen obtained from the combined effluent does not contain carbon monoxide or carbon dioxide and therefore can be fully utilised to react with field carbon dioxide.
As mentioned hereinabove it is preferred to use both carbon dioxide and carbon monoxide to synthesise the oxygenates, while it is also preferred to combine the carbon dioxide ex. field with a synthesis that is rich in hydrogen. By providing at least part of the hydrogen obtained from the combined effluent the molar ratio of the hydrogen, carbon dioxide and optionally carbon monoxide provided to the oxygenate synthesis zone of step e) can be increased. It is particularly preferred in a process, wherein the synthesis gas produced in step (c) is hydrogen deficient, in particular a synthesis gas having a molar ratio of hydrogen, carbon dioxide and optionally carbon monoxide in the range of from 1.0 to 1.9, more preferably of from 1.3 to 1.8, wherein the molar ratio is defined hereinabove. Such a synthesis gas is typically produced with little or even no steam added to the process. As a result the water consumption of the synthesis gas process is reduced, and thus the water consumption of the integrated process is reduced.
The synthesis gas is combined with hydrogen, for example obtained from the combined effluent. For example, a synthesis gas comprising mainly hydrogen and carbon monoxide at a molar ratio of 1.6, may be combined with hydrogen, for example obtained from step (d) of the process according to the invention to form a hydrogen-enriched synthesis gas, which can be mixed with at least part of the carbon dioxide ex. field. Preferably, sufficient field carbon dioxide is added to the synthesis gas to provide a carbon dioxide concentration in the range of from 0.1 to 25 mol %, preferably 3 to 15 mol %, more preferably of from 4 to 10 mol %, based on the total number of moles of hydrogen, carbon monoxide and carbon dioxide in the mixture.
As mentioned, preferably, a synthesis gas is used comprising little or no carbon dioxide. The carbon dioxide ex. field preferably comprises little or no inerts like Ar, or N2. When using a synthesis gas with little or no carbon dioxide, more field carbon dioxide can be added and less inerts are introduced to the oxygenate synthesis zone. Less waste carbon dioxide is thus produced, which would otherwise need to be sequestrated or captured and stored.
By using hydrogen obtained from combined effluent to produce at least part of the oxygenate feed to the OTO zone, the hydrogen produced during the cracking step or OTO is no longer combusted as fuel in the ethane furnace, but rather used to produce valuable oxygenates. This hydrogen is obtained as co-product and does not add additional carbon dioxide on top of what is required for the main reaction product ethylene. In addition, the hydrogen obtained from the combined effluent does not comprise significant amounts of inerts such as N2, Ar or CH4. These inerts may typically be present in the natural gas or purified oxygen provided to produce synthesis gas for methanol production.
By combining at least part of the cracking zone effluent and at least part of the OTO zone effluent to a combined effluent, at least part of any olefins other than ethylene obtained in step (d) and step (f) are also combined in one stream. The cracker zone effluent obtained from the cracking zone of step (d) comprises predominantly ethylene, but may also comprise up to 2 wt % propylene, based on the total weight of the ethylene in the cracking zone effluent. This amount of propylene is not economically recoverable, however by combining the olefins obtained from the cracking zone in step (d) and the olefins obtained from the OTO zone in step (f), i.e. combining at least part of the cracking zone effluent and at least part of the OTO zone effluent to a combined effluent, a combined effluent is obtained that comprises in the range of 10 to 40% of propylene, based on the total hydrocarbon content of the combined effluent. The high propylene content in the combined effluent is caused by the high propylene content in the OTO zone effluent. OTO processes produce a mix of olefins, comprising in the range of from 5 to 80 wt % of ethylene and in the range of from 10 to 80 wt % of propylene, based on the total hydrocarbon content of the OTO zone effluent. By combining the effluent of the cracking zone and the OTO zone, it is possible to also economically recover the propylene in the cracking zone effluent. The propylene can be used as a feedstock to a polypropylene production process, optionally after being treated to remove impurities. Polypropylene production processes are well known in the art.
Besides olefins and hydrogen, the OTO process also produces small amounts of alkanes, in particular ethane, propane and butane. A further synergy of the integrated process can be obtained by providing any ethane present in the effluent of the OTO zone to the cracking zone. The ethane can then be cracked to ethylene and hydrogen in the cracking zone, thus providing additional ethylene and hydrogen. The hydrogen may subsequently be used to synthesise oxygenates.
Part of the ethylene produced in the process according to the invention can be used as a feedstock for several other processes, including the synthesis of ethylene oxide, mono-ethylene-glycol and styrene monomer.
It has now also been found that it is possible to integrate the production of these products into the process according to the invention to obtain further synergy.
A further integration may be achieved by converting at least part of the ethylene produced in step (d), step (f) or, preferably, both steps (d) and (f) with benzene into ethyl benzene and converting at least part of the ethyl benzene to styrene monomer and at least hydrogen.
Each of the mentioned conversion steps hereinabove are well known in the art. Any suitable process may be used. Ethyl benzene is typically produced by reacting ethylene and benzene in the presence of an acid catalyst. Reference is for example made to made to Kniel et al., Ethylene, Keystone to the petrochemical industry, Marcel Dekker, Inc, New York, 1980, in particular section 3.4.1, page 24 to 25. While the styrene is produced by the catalytic dehydrogenation of ethyl benzene in the presence of a suitable catalyst, examples of suitable catalysts include but are not limited to dehydrogenation catalysts based on iron (III) oxide.
By integrating the process according to the invention with the production of styrene monomer, as described above, further hydrogen is produced next to the desired products. Preferably, this hydrogen is separated from the other reaction products. As mentioned hereinabove, additional hydrogen, for instance produced in step (d) of the process, can be used to provide additional hydrogen to the oxygenate synthesis zone. Also, the hydrogen obtained from the conversion of ethylene, via ethylbenzene, to styrene monomer may preferably be used as additional hydrogen to the oxygenate synthesis zone.
By providing additional hydrogen obtained from the conversion of ethylene, via ethylbenzene, into styrene to synthesise oxygenates, more field carbon dioxide can be provided to the oxygenate synthesis zone.
The produced styrene monomer may be used to produce polystyrene.
Carbon dioxide other than field carbon dioxide may also be provided to the oxygenate synthesis zone, in particular in case the amount of carbon dioxide ex. field is insufficient to satisfy the full carbon dioxide demand of the oxygenate synthesis. As mentioned hereinabove some additional carbon dioxide may for instance be comprised in synthesis gas provided to the oxygenate synthesis as part of the feed containing hydrogen. Other suitable sources of carbon dioxide may include:

    
    
        a source comprising carbon dioxide obtained from a carbon dioxide-comprising flue gas stream, in particular a flue gas obtained from the integrated process according to the invention or optionally an oxygen purification unit or synthesis gas production process. The oxidation of ethylene to ethylene oxide requires large amounts of purified oxygen. Preferably, the flue gas is first concentrated to increase the carbon dioxide concentration.
        a source comprising the flue gas obtained from an oxidative de-coking of an ethane cracking furnace, typically one of the ethane cracking furnaces used for producing olefins in step (d). In case the oxidative de-coking of the furnace is done using pure oxygen or pure oxygen diluted in carbon dioxide instead of air, an essentially pure stream of carbon dioxide can be produced, which is especially suitable to be used in the synthesis of oxygenates. Although, it is required to first produce pure oxygen, there is no need to post-treat the flue gas in order to capture the carbon dioxide. Also the decoking to the OTO catalyst during regeneration of the catalyst can be performed in a similar way to provide a suitable carbon dioxide comprising stream.
    
    


A particularly preferred source of carbon dioxide may be obtained by converting ethylene to ethylene oxide or MEG. Ethylene oxide and MEG are valuable products, whereby MEG has the further advantage that it is liquid and therefore can be stored and transported relatively easy.
In the process according to the invention at least part of obtained ethylene from the ethane cracking step (d) and/or the OTO step (f), preferably both steps (d) and (f) may be oxidised to ethylene oxide by providing at least part of the ethylene with a feed containing oxygen to an ethylene oxidation zone, further referred to as EO zone.
Preferably, the ethylene oxide is further converted to mono-ethylene-glycol (MEG). MEG is a liquid and therefore can be transported and stored more conveniently than ethylene oxide. Preferably, the EO zone is part of a larger mono-ethylene-glycol synthesis zone, i.e. a second oxygenate synthesis zone, further referred to as MEG zone. Preferably, the MEG zone then comprises a first section comprising the EO zone and a second ethylene oxide hydrolysis section. The MEG is synthesised by providing the ethylene oxide with a feed containing water to the ethylene oxide hydrolysis zone and converting the ethylene oxide to MEG. Optionally, the ethylene oxide is first reacted with carbon dioxide to from ethylene carbonate, which is subsequently hydrolysed to obtain MEG and carbon dioxide, reference herein is made to for instance US2008139853, incorporated by reference.
A by-product of the ethylene oxide process is carbon dioxide. During the reaction of ethylene with oxygen to ethylene oxide, carbon dioxide is formed as part of the ethylene is fully oxidised to carbon dioxide. The carbon dioxide is typically obtained as part of an EO zone effluent, which also comprises ethylene oxide. This is waste carbon dioxide and needs to be sequestered or otherwise captured and stored.
By providing part of the carbon dioxide obtained as part of an EO zone effluent to the oxygenate synthesis zone of step (e) the need to sequester or otherwise capture or store this waste carbon dioxide is reduced.
Preferably, the carbon dioxide is separated from the OE zone effluent to obtain a separate carbon dioxide comprising stream, further also referred to a carbon dioxide ex. EO. Preferably, the EO zone effluent is further treated to convert the ethylene oxide into MEG in a MEG zone. From the MEG zone, a MEG zone effluent is obtained, comprising MEG and carbon dioxide. Suitably, the carbon dioxide can be separated from the MEG zone effluent by cooling the MEG zone effluent to a temperature below the boiling point of MEG, this carbon dioxide is also referred to as carbon dioxide ex. MEG. As no additional carbon dioxide is produced by converting ethylene oxide into MEG the carbon dioxide ex. MEG is the same as the carbon dioxide ex. EO. By reusing the carbon dioxide to synthesize oxygenates, the carbon dioxide penalty for producing EO is reduced. A further advantage is that the stream comprising carbon dioxide obtained from the EO or MEG zone comprises predominantly carbon dioxide. Preferably the stream comprises in the range of 80 to 100 mol % of carbon dioxide and steam, based on the total amount of moles in the stream. More preferably, the stream comprising carbon dioxide comprises essentially only carbon dioxide and, optionally, steam. Such a stream is particularly suitable to be used in an oxygenate synthesis process as it does not introduce significant amounts of inerts, e.g. CH4, N2 and Ar, to the oxygenate synthesis zone. Should, however, the stream comprising carbon dioxide comprise significant amounts of other, undesired, compounds, e.g. ethylene oxide, the stream is preferably treated to remove such compounds prior to being introduced into the oxygenate synthesis zone. Another advantage of the integration with a MEG synthesis is that next to MEG, minor amounts of other oxygenates may produced in the MEG zone by the process for producing MEG, such as for instance di-ethylene-glycol. These oxygenates may suitably be separated from the obtained MEG zone effluent and provided to the OTO zone as part of the oxygenate feed.
In the present invention an ethane-comprising feed is cracked in step (d) in a cracking zone under cracking conditions to produce at least olefins and hydrogen.
Additionally, small amounts of propylene are formed. Other by-products may be formed such as butylene, butadiene, ethyne, propyne and benzene. The cracking process is performed at elevated temperatures, preferably in the range of from 650 to 1000° C., more preferably from 750 to 950° C. Typically, the cracking is performed in the presence of water (steam) as a diluent. The ethane conversion is typically in the range of from 40 to 75 mol %, based on the total number of moles of ethane provided to the cracking zone. Preferably, the un-cracked ethane is recycled back to the cracking zone. Ethane cracking processes are well known to the skilled person and need no further explanation. Reference is for instance made to Kniel et al., Ethylene, Keystone to the petrochemical industry, Marcel Dekker, Inc, New York, 1980, in particular chapters 6 and 7.
In the present invention an oxygenate feedstock is converted in step (f) in an oxygenate-to-olefins process, in which an oxygenate feedstock is contacted in an OTO zone with an oxygenate conversion catalyst under oxygenate conversion conditions, to obtain a conversion effluent comprising lower olefins. In the OTO zone, at least part of the feed is converted into a product containing one or more olefins, preferably including light olefins, in particular ethylene and/or propylene.
Examples of oxygenates that can be used in the oxygenate feedstock of step f) of the process include alcohols, such as methanol, ethanol, isopropanol, ethylene glycol, propylene glycol; ketones, such as acetone and methylethylketone; aldehydes, such as formaldehyde, acetaldehyde and propionaldehyde; ethers, such as dimethylether, diethylether, methylethylether, tetrahydrofuran and dioxane; epoxides such as ethylene oxide and propylene oxide; and acids, such as acetic acid, propionic acid, formic acid and butyric acid. Further examples are dialkyl carbonates such as dimethyl carbonate or alkyl esters of carboxylic acids such as methyl formate. Of these examples, alcohols and ethers are preferred.
Examples of preferred oxygenates include alcohols, such as methanol, ethanol, isopropanol, ethylene glycol, propylene glycol; and dialkyl ethers, such as dimethylether, diethylether, methylethylether. Cyclic ethers such as tetrahydrofuran and dioxane, are also suitable.
The oxygenate used in the process according to the invention is preferably an oxygenate which comprises at least one oxygen-bonded alkyl group. The alkyl group preferably is a C1-C4 alkyl group, i.e. comprises 1 to 4 carbon atoms; more preferably the alkyl group comprises 1 or 2 carbon atoms and most preferably one carbon atom. The oxygenate can comprise one or more of such oxygen-bonded C1-C4 alkyl groups. Preferably, the oxygenate comprises one or two oxygen-bonded C1-C4 alkyl groups.
More preferably an oxygenate is used having at least one C1 or C2 alkyl group, still more preferably at least one C1 alkyl group.
Preferably the oxygenate is chosen from the group of alkanols and dialkyl ethers consisting of dimethylether, diethylether, methylethylether, methanol, ethanol, isopropanol, and mixtures thereof.
Most preferably the oxygenate is methanol or dimethylether, or a mixture thereof.
Preferably the oxygenate feedstock comprises at least 50 wt % of oxygenate, in particular methanol and/or dimethylether, based on total hydrocarbons, more preferably at least 80 wt %, most preferably at least 90 wt %.
The oxygenate feedstock can be obtained from a prereactor, which converts methanol at least partially into dimethylether. In this way, water may be removed by distillation and so less water is present in the process of converting oxygenate to olefins, which has advantages for the process design and lowers the severity of hydrothermal conditions the catalyst is exposed to.
The oxygenate feedstock can comprise an amount of diluents, such as water or steam.
A variety of OTO processes are known for converting oxygenates such as for instance methanol or dimethylether to an olefin-containing product, as already referred to above. One such process is described in WO-A 2006/020083, incorporated herein by reference, in particular in paragraphs [0116]-[0135]. Processes integrating the production of oxygenates from synthesis gas and their conversion to light olefins are described in US2007/0203380A1 and US2007/0155999A1.
Catalysts as described in WO A 2006/020083 are suitable for converting the oxygenate feedstock in step (f) of the present invention. Such catalysts preferably include molecular sieve catalyst compositions. Suitable molecular sieves are silicoaluminophosphates (SAPO), such as SAPO-17, -18, -34, -35, -44, but also SAPO-5, -8, -11, -20, -31, -36, -37, -40, -41, -42, -47 and -56.
Alternatively, the conversion of the oxygenate feedstock may be accomplished by the use of an aluminosilicate catalyst, in particular a zeolite. Suitable catalysts include those containing a zeolite of the ZSM group, in particular of the MFI type, such as ZSM-5, the MTT type, such as ZSM-23, the TON type, such as ZSM-22, the MEL type, such as ZSM-11, the FER type. Other suitable zeolites are for example zeolites of the STF-type, such as SSZ-35, the SFF type, such as SSZ-44 and the EU-2 type, such as ZSM-48. Aluminosilicate catalysts are preferred when an olefinic co-feed is fed to the oxygenate conversion zone together with oxygenate, for increased production of ethylene and propylene.
The reaction conditions of the oxygenate conversion include those that are mentioned in WO-A 2006/020083. Hence, a reaction temperature of 200 to 1000° C., preferably from 250 to 750° C., and a pressure from 0.1 kPa (1 mbar) to 5 MPa (50 bar), preferably from 100 kPa (1 bar) to 1.5 MPa (15 bar), are suitable reaction conditions.
A specially preferred OTO process for use in step (f) of the present invention will now be described. This process provides particularly high conversion of oxygenate feed and a recycle co-feed to ethylene and propylene. Reference is made in this regard also to WO2007/135052, WO2009/065848, WO2009/065875, WO2009/065870, WO2009/065855, WO2009/065877, in which processes a catalyst comprising an aluminosilicate or zeolite having one-dimensional 10-membered ring channels, and an olefinic co-feed and/or recycle feed is employed.
In this process, the oxygenate-conversion catalyst comprises one or more zeolites having one-dimensional 10-membered ring channels, which are not intersected by other channels, preferably at least 50% wt of such zeolites based on total zeolites in the catalyst. Preferred examples are zeolites of the MTT and/or TON type. In a particularly preferred embodiment the catalyst comprises in addition to one or more one-dimensional zeolites having 10-membered ring channels, such as of the MTT and/or TON type, a more-dimensional zeolite, in particular of the MFI type, more in particular ZSM-5, or of the MEL type, such as zeolite ZSM-11. Such further zeolite (molecular sieve) can have a beneficial effect on the stability of the catalyst in the course of the OTO process and under hydrothermal conditions. The second molecular sieve having more-dimensional channels has intersecting channels in at least two directions. So, for example, the channel structure is formed of substantially parallel channels in a first direction, and substantially parallel channels in a second direction, wherein channels in the first and second directions intersect. Intersections with a further channel type are also possible. Preferably the channels in at least one of the directions are 10-membered ring channels. A preferred MFI-type zeolite has a Silica-to-Alumina ratio SAR of at least 60, preferably at least 80, more preferably at least 100, even more preferably at least 150. The oxygenate conversion catalyst can comprise at least 1 wt %, based on total molecular sieve in the oxygenate conversion catalyst, of the second molecular sieve having more-dimensional channels, preferably at least 5 wt %, more preferably at least 8 wt %, and furthermore can comprise less than 35 wt % of the further molecular sieve, in certain embodiments less than 20 wt %, or less than 18 wt %, such as less than 15 wt %.
Especially when the oxygenate conversion is carried out over a catalyst containing MTT or TON type aluminosilicates, it may be advantageous to add an olefin-containing co-feed together with the oxygenate feed (such as dimethylether-rich or methanol-rich) feed to the OTO zone when the latter feed is introduced into this zone. It has been found that the catalytic conversion of the oxygenates, in particular methanol and DME, to ethylene and propylene is enhanced when an olefin is present in the contact between methanol and/or dimethylether and the catalyst. Therefore, suitably, an olefinic co-feed is added to the reaction zone together with the oxygenate feedstock.
In special embodiments, at least 70 wt % of the olefinic co-feed, during normal operation, is formed by a recycle stream of a C3+ or C4+ olefinic fraction from the OTO conversion effluent, preferably at least 90 wt %, more preferably at least 99 wt %, and most preferably the olefinic co-feed is during normal operation formed by such recycle stream. In one embodiment the olefinic co-feed can comprise at least 50 wt % of C4 olefins, and at least a total of 70 wt % of C4 hydrocarbon species. It can also comprise propylene. The OTO conversion effluent can comprise 10 wt % or less, preferably 5 wt % or less, more preferably 1 wt % or less, of C6-C8 aromatics, based on total hydrocarbons in the effluent. At least one of the olefinic co-feed, and the recycle stream, can in particular comprise less than 20 wt % of C5+ olefins, preferably less than 10 wt % of C5+ olefins, based on total hydrocarbons in the olefinic co-feed.
In order to maximize production of ethylene and propylene, it is desirable to maximize the recycle of C4 olefins. In a stand-alone process, i.e. without integration with a cracker, there is a limit to the maximum recycle of a C4 fraction from the OTO effluent. A certain part thereof, such as between 1 and 5 wt %, needs to be withdrawn as purge, since otherwise saturated C4's (butane) would build up which are substantially not converted under the OTO reaction conditions.
In the preferred process, optimum light olefins yield are obtained when the OTO conversion is conducted at a temperature of more than 450° C., preferably at a temperature of 460° C. or higher, more preferably at a temperature of 480° C. or higher, in particular at 500° C. or higher, more in particular 550° C. or higher, or 570° C. or higher. The temperature will typically less than 700° C., or less than 650° C. The pressure will typically be between 0.5 and 15 bar, in particular between 1 and 5 bar.
In a special embodiment, the oxygenate conversion catalyst comprises more than 50 wt %, preferably at least 65 wt %, based on total molecular sieve in the oxygenate conversion catalyst, of the one-dimensional molecular sieve having 10-membered ring channels.
In one embodiment, molecular sieves in the hydrogen form are used in the oxygenate conversion catalyst, e.g., HZSM-22, HZSM-23, and HZSM-48, HZSM-5. Preferably at least 50% w/w, more preferably at least 90% w/w, still more preferably at least 95% w/w and most preferably 100% of the total amount of molecular sieve used is in the hydrogen form. When the molecular sieves are prepared in the presence of organic cations the molecular sieve may be activated by heating in an inert or oxidative atmosphere to remove organic cations, for example, by heating at a temperature over 500° C. for 1 hour or more. The zeolite is typically obtained in the sodium or potassium form. The hydrogen form can then be obtained by an ion exchange procedure with ammonium salts followed by another heat treatment, for example in an inert or oxidative atmosphere at a temperature over 500° C. for 1 hour or more. The molecular sieves obtained after ion-exchange are also referred to as being in the ammonium form.
The molecular sieve can be used as such or in a formulation, such as in a mixture or combination with a so-called binder material and/or a filler material, and optionally also with an active matrix component. Other components can also be present in the formulation. If one or more molecular sieves are used as such, in particular when no binder, filler, or active matrix material is used, the molecular sieve itself is/are referred to as oxygenate conversion catalyst. In a formulation, the molecular sieve in combination with the other components of the mixture such as binder and/or filler material is/are referred to as oxygenate conversion catalyst.
It is desirable to provide a catalyst having good mechanical or crush strength, because in an industrial environment the catalyst is often subjected to rough handling, which tends to break down the catalyst into powder-like material. The latter causes problems in the processing. Preferably the molecular sieve is therefore incorporated in a binder material. Examples of suitable materials in a formulation include active and inactive materials and synthetic or naturally occurring zeolites as well as inorganic materials such as clays, silica, alumina, silica-alumina, titania, zirconia and aluminosilicate. For present purposes, inactive materials of a low acidity, such as silica, are preferred because they may prevent unwanted side reactions which may take place in case a more acidic material, such as alumina or silica-alumina is used.
Typically the oxygenate conversion catalyst deactivates in the course of the process. Conventional catalyst regeneration techniques can be employed. The catalyst particles used in the process of the present invention can have any shape known to the skilled person to be suitable for this purpose, for it can be present in the form of spray dried catalyst particles, spheres, tablets, rings, extrudates, etc. Extruded catalysts can be applied in various shapes, such as, cylinders and trilobes. If desired, spent oxygenate conversion catalyst can be regenerated and recycled to the process of the invention. Spray-dried particles allowing use in a fluidized bed or riser reactor system are preferred. Spherical particles are normally obtained by spray drying. Preferably the average particle size is in the range of 1-200 μm, preferably 50-100 μm.
The preferred embodiment of step (f) described hereinabove is preferably performed in an OTO zone comprising a fluidized bed or moving bed, e.g. a fast fluidized bed or a riser reactor system, although in general for an OTO process, in particular for an MTP process, also a fixed bed reactor or a tubular reactor can be used. Serial reactor systems can be employed.
In one embodiment, the OTO zone comprises a plurality of sequential reaction sections. Oxygenate can be added to at least two of the sequential reaction sections.
When multiple reaction zones are employed, an olefinic co-feed is advantageously added to the part of the dimethylether-rich feed that is passed to the first reaction zone.
The preferred molar ratio of oxygenate in the oxygenate feedstock to olefin in the olefinic co-feed provided to the OTO conversion zone depends on the specific oxygenate used and the number of reactive oxygen-bonded alkyl groups therein. Preferably the molar ratio of oxygenate to olefin in the total feed lies in the range of 20:1 to 1:10, more preferably in the range of 18:1 to 1:5 and still more preferably in the range of 15:1 to 1:3.
A diluent can also be fed to the OTO zone, mixed with the oxygenate and/or co-feed if present, or separately. A preferred diluent is steam, although other inert diluents can be used as well. In one embodiment, the molar ratio of oxygenate to diluent is between 10:1 and 1:10, preferably between 4:1 and 1:2, most preferably between 3:1 and 1:1, such as 1.5:1, in particular when the oxygenate is methanol and the diluent is water (steam).
The feedstock comprising methane, ethane and carbon dioxide may be any feedstock comprising methane, ethane and carbon dioxide. Preferably, the feedstock is natural gas or associated gas. Reference herein to associated gas is to C1 to C5 hydrocarbons co-produced with the production of oil. Preferably the feedstock comprises in the range of from 0.1 to 70 mol % of carbon dioxide, more preferably from 0.5 to 35 mol %, even more preferably from 1 to 30 mol %, based on the total content of the feedstock. The feedstock comprising methane, ethane and carbon dioxide may further comprise higher hydrocarbons, including but not limited to, propane, butanes and pentanes.
Preferably, the feed comprising methane and ethane comprises in the range of from 1 to 20 mol % of ethane, based on the total feed.
The methane-comprising feed, preferably, comprises in the range of from 50 to 100 mol % of methane, more preferably 80 to 99 mol % of methane, based on the total number of moles in the methane-comprising feed.
The ethane-comprising feed, preferably, comprises in the range of from 50 to 100 mol % of ethane, more preferably 80 to 99 mol % of ethane, based on the total number of moles in the ethane-comprising feed.
The carbon dioxide-comprising feed (or carbon dioxide ex. field), preferably, comprises in the range of from 90 to 100 mol % of carbon dioxide, more preferably 97 to 99 mol % of carbon dioxide, based on the total number of moles in the carbon dioxide-comprising feed.
The oxygenates provided to step (f) of the process according to the invention may be any oxygenates preferably methanol and/or DME. The oxygenate feedstock comprises at least oxygenates obtained in step (e). The oxygenates may further comprise oxygenates, such as for example other alcohols, other ethers, aldehydes, ketones and esters. Preferably, the oxygenates are provided together with water as a diluent. An oxygenate feedstock provided to the process may also comprise compounds other than water and oxygenates.
In one embodiment, the oxygenate is obtained as a reaction product of synthesis gas. Synthesis gas can for example be generated from fossil fuels, such as from natural gas or oil, or from the gasification of coal. Suitable processes for this purpose are for example discussed in Industrial Organic Chemistry, Klaus Weissermehl and Hans-Jürgen Arpe, 3rd edition, Wiley, 1997, pages 13-28. This book also describes the manufacture of methanol from synthesis gas on pages 28-30.
In another embodiment the oxygenate is obtained from biomaterials, such as through fermentation. For example by a process as described in DE-A-10043644.
Preferably, at least part of the oxygenate feed is obtained by converting methane into synthesis gas and providing the synthesis gas to a oxygenate synthesis zone to synthesise oxygenates. The methane is preferably obtained from natural gas or associated gas, more preferably the same natural gas or associated gas, from which the light paraffin feedstock for the cracker is obtained.
The oxygenates may be provided directly from one or more oxygenate synthesis zones, however, it may also be provided from a central oxygenate storage facility.
The olefinic co-feed optionally provided together with the oxygenate feedstock to the OTO conversion zone may contain one olefin or a mixture of olefins. Apart from olefins, the olefinic co-feed may contain other hydrocarbon compounds, such as for example paraffinic, alkylaromatic, aromatic compounds or a mixture thereof. Preferably the olefinic co-feed comprises an olefinic fraction of more than 20 wt %, more preferably more than 25 wt %, still more preferably more than 50 wt %, which olefinic fraction consists of olefin(s). The olefinic co-feed can consist essentially of olefin(s).
Any non-olefinic compounds in the olefinic co-feed are preferably paraffinic compounds. If the olefinic co-feed contains any non-olefinic hydrocarbon, these are preferably paraffinic compounds. Such paraffinic compounds are preferably present in an amount in the range from 0 to 80 wt %, more preferably in the range from 0 to 75 wt %, still more preferably in the range from 0 to 50 wt %.
By an unsaturate is understood an organic compound containing at least two carbon atoms connected by a double or triple bond. By an olefin is understood an organic compound containing at least two carbon atoms connected by a double bond. The olefin can be a mono-olefin, having one double bond, or a poly-olefin, having two or more double bonds. Preferably olefins present in an olefinic co-feed are mono-olefins. C4 olefins, also referred to as butenes (1-butene, 2-butene, iso-butene, and/or butadiene), in particular C4 mono-olefins, are preferred components in the olefinic co-feed.
Preferred olefins have in the range from 2 to 12, preferably in the range from 3 to 10, and more preferably in the range from 4 to 8 carbon atoms.
Examples of suitable olefins that may be contained in the olefinic co-feed include ethene, propene, butene (one or more of 1-butene, 2-butene, and/or iso-butene (2-methyl-1-propene)), pentene (one or more of 1-pentene, 2-pentene, 2-methyl-1-butene, 2-methyl-2-butene, 3-methyl-1-butene, and/or cyclopentene), hexene (one or more of 1-hexene, 2-hexene, 3-hexene, 2-methyl-1-pentene, 2-methyl-2-pentene, 3-methyl-1-pentene, 3-methyl-2-pentene, 4-methyl-1-pentene, 4-methyl-2-pentene, 2,3-dimethyl-1-butene, 2,3-dimethyl-2-butene, 3,3-dimethyl-1-butene, methylcyclopentene and/or cyclohexene), heptenes, octenes, nonenes and decenes. The preference for specific olefins in the olefinic co-feed may depend on the purpose of the process, such as preferred production of ethylene or propylene.
In a preferred embodiment the olefinic co-feed preferably contains olefins having 4 or more carbon atoms (i.e. C4+ olefins), such as butenes, pentenes, hexenes and heptenes. More preferably the olefinic fraction of the olefinic co-feed comprises at least 50 wt % of butenes and/or pentenes, even more preferably at least 50% wt of butenes, and most preferably at least 90 wt % of butenes. The butene may be 1-, 2-, or iso-butene, or a mixture of two or more thereof.
Preferably, at least part of the oxygenate feed is obtained by converting methane into synthesis gas and providing the synthesis gas to a oxygenate synthesis zone to synthesise oxygenates. The methane is preferably obtained from natural gas or associated gas, more preferably the same natural gas or associated gas, from which the ethane-comprising feed was obtained.
The benzene used to convert ethylene into ethyl benzene may be any benzene available. The benzene may be benzene produced in step (a) of the process according to the invention. As disclosed in U.S. Pat. No. 6,677,496, ethane cracking processes typically produce up to 0.6 wt % of benzene, based on the total ethane feed. However, the benzene may also be obtained from any other source.
Preferably, the benzene is produced from higher hydrocarbons such as propane and butane, more preferably propane and butane obtained as condensate or LPG from the natural gas or associated gas which served as the feedstock comprising methane, ethane and carbon dioxide.
In another aspect, the invention provides an integrated system for producing olefins, which system comprises:

    
    
        a) a separation system, having at least an inlet for a feed comprising methane, ethane and carbon dioxide, and an outlet for a methane-comprising feed, an outlet for an ethane-comprising feed and an outlet for a carbon dioxide-comprising feed;
        b) a partial oxidation system arranged to partially oxidise the methane-comprising feed to a synthesis gas, having an inlet for a methane-comprising feed to receive the methane-comprising feed from the separation system, an inlet for oxygen, and an outlet for synthesis gas;
        c) a steam cracking system having one or more inlets for an ethane-comprising feed to receive the methane-comprising feed from the separation system and an inlet for steam, and an outlet for a cracker effluent comprising olefins;
        d) an oxygenate-to-olefins conversion system, having one or more inlets for receiving an oxygenate feedstock, and comprising a reaction zone for contacting the oxygenate feedstock with an oxygenate conversion catalyst under oxygenate conversion conditions, and an outlet for an oxygenate-to-olefins effluent comprising olefins, including at least ethylene;
        e) a work-up system arranged to receive at least part of the cracker effluent and at least part of the oxygenate-to-olefins effluent to obtain a combined effluent, the work-up section comprising a separation system, an outlet for ethylene and an outlet for hydrogen;
        f) an oxygenate synthesis system having one or more inlets for synthesis gas to receive the synthesis gas from the partial oxidation system, an inlet for hydrogen and an inlet for a carbon dioxide-comprising feed to receive the carbon dioxide-comprising feed from the separation system, and an outlet for an oxygenate feedstock; and

means for providing the oxygenate feedstock from the outlet for oxygenate feedstock of the oxygenate synthesis system to the oxygenate feedstock inlet of the oxygenate-to-olefins conversion system and means for providing hydrogen from the outlet for hydrogen of the work-up section to the inlet for hydrogen of the oxygenate synthesis system.

    
    


The inlet for hydrogen of the oxygenate synthesis system may be the same as the inlet for synthesis gas.
In FIG. 1, a schematic representation is given of an embodiment of integrated system for producing olefins according to the invention. In the system of FIG. 1, a ethane-comprising feed and steam are provided via conduits 1 and 3 respectively to steam cracking system 5, comprising a cracking zone for steam cracking ethane to ethylene. From steam cracking system 5, a cracking zone effluent is retrieved via conduit 7.
In FIG. 1, also a synthesis gas and a carbon dioxide-comprising are provided via conduits 9 and 10, respectively, to oxygenate synthesis system 11, comprising an oxygenate synthesis zone for synthesizing oxygenates from hydrogen and at least one of carbon monoxide and carbon dioxide. From oxygenate synthesis system 11, an oxygenate feedstock is retrieved via conduit 13. The oxygenate feedstock is provided to oxygenate-to-olefins conversion system 15, comprising an OTO zone for converting oxygenates to lower olefins, e.g. ethylene and propylene. Optionally, an olefinic co-feed (not shown) is provided to oxygenate-to-olefins conversion system 15 together with the oxygenate feedstock. From oxygenate-to-olefins conversion system 13, an OTO zone effluent is retrieved via conduit 17.
The cracking zone effluent and the OTO zone effluent are combined to form combined effluent in conduit 19 and provided to work-up section 21. Work-up section 21 comprises at least a separation system to separate ethylene from the combined effluent. In addition, hydrogen is separated from the combined effluent and provided to oxygenate synthesis system 11 via conduit 23. The ethylene is retrieved separately from workup section 21 and provided via conduit 25 to ethylene oxidation system 27, which comprises an oxygenation zone for oxidising ethylene to ethylene oxide. Oxygen is provided to ethylene oxidation system 27 via conduit 29. From ethylene oxide system 27, ethylene oxide is retrieved via conduit 31 and provided to ethylene oxide hydrolysis system 33, which comprises an ethylene oxide hydrolysis zone, wherein ethylene oxide is hydrolysed to MEG. Water is provided to ethylene oxide hydrolysis system 33 via conduit 35. Ethylene oxidation system 27 and ethylene oxide hydrolysis system 33 are comprised in MEG synthesis system 37. From MEG synthesis system 37, a MEG-comprising effluent is retrieved via conduit 39 and carbon dioxide via conduit 41. In FIG. 1, a feed comprising methane, ethane and carbon dioxide is provided via conduit 101 to separation system 103. In separation system 103, the feed comprising methane, ethane and carbon dioxide is separated into at least a methane-comprising feed, an ethane-comprising feed and a carbon dioxide-comprising feed. From separation system 103, the ethane comprising feed is provided via conduit 1 to steam cracking system 5. The methane-comprising feed is provided via conduit 107 to partial oxidation system 109 to be partially oxidised to form synthesis gas. Oxygen is provided via conduit 111 to partial oxidation system 109. From partial oxidation system 109, the synthesis gas is retrieved via conduit 9 and provided to oxygenate synthesis system 11. The carbon dioxide-comprising feed is provided from separation system 103 to oxygenate synthesis system 11 via conduit 10.
EXAMPLES
The invention is illustrated by the following non-limiting calculated examples.
Example 1
In the Examples, several options of implementing the present invention are compared with comparative examples, by means of model calculations. As basis for Examples 1a to g, a model integrated OTO/ethane cracker process was taken. In Table 1, an overview is provided of the feed input and the calculated products.
Calculations were done using a Spyro based model for modelling of cracker operation combined with a proprietary model for modelling the OTO conversion. The key input to the models was as follows:
Cracking:
Steam to ethane ratio is 0.35 wt %. USC coil is used for the Spyro calculations. Calculated at a coil outlet pressure of 1.77 bar absolute, at 65% ethane conversion and a residence time of 0.24 seconds.
OTO Conversion:
MeOH 5012 t/d is fed to the OTO reactor together with 1384 t/d of recycled and superheated steam and 1775 t/d of recycled C4 stream. The model was calibrated on small-scale experiments conducted to determine product distributions for single-pass OTO conversions. Therein, all components that were fed to the OTO reactor have been evaporated and heated such that the temperature in the reactor is controlled between 550-600° C. The pressure in the reactor is 2 bar absolute. The OTO catalyst is fluidized in the reaction medium under the condition that the weight hourly space velocity (WHSV) is 4-10 h−1, whereby WHSV is defined as the total weight of the feed flow over the catalyst weight per hour. The following catalyst was used: Composition and preparation: 32 wt % ZSM-23 SAR 46, 8 wt % ZSM-5 SAR 280, 36 wt % kaolin, 24 wt % silica sol, and, after calcination of the ammonium form of the spray dried particle, 1.5 wt % P was introduced by H3PO4 impregnation. The catalyst was again calcined at 550° C. The steam and C4 recycle streams are excluded from the product composition tables.
The methanol provided to the OTO process (approximately 5000 t/d, see Table 1) is synthesised using at least part of the field carbon dioxide.
As feedstock comprising methane, ethane and carbon dioxide, a natural gas was used having the following composition: 89.8 mol % CH4, 5.4 mol % C2H6, 4.4 mol % N2, 0.4 mol % CO2, and 0.1 mol % Ar, based on the total number of moles in the natural gas stream.
The ethane-comprising feed is obtained by separating the ethane from the natural gas. The remainder of the natural gas is used as the methane-comprising feed. The composition of the methane-comprising feed is: 94.3 mol % CH4, 0.6 mol % C2H6, 4.6 mol % N2, 0.4 mol % CO2, and 0.1 mol % Ar, based on the total number of moles in the methane-comprising feed.
Additional carbon dioxide is added to simulate a carbon dioxide stream separated from the natural gas.
Optionally, additional hydrogen obtained from the combined effluent is added.
Both hydrogen ex. combined effluent and carbon dioxide ex. field are taken as 99.9+% pure.
The yields of methanol are calculated by an Aspen model. To keep the inert concentration of about 40 wt % in the synthesis gas recycle, the amount of purge stream from the recycle is adjusted.
The methane-comprising feed is converted to synthesis gas using different processes resulting in the following synthesis gasses:
Synthesis gas from a non-catalytic partial oxidation of natural gas (Shell gasification process). The SGP syngas comprises 61.2 mol % hydrogen, 34.0 mol % carbon monoxide, 2.1 mol % carbon dioxide and 2.5 mol % inerts (N2, Ar and CH4), based on the total number of moles in the SGP syngas.
Synthesis gas from an auto-thermal reforming of natural gas (ATR). The ATR syngas comprises 65.5 mol % hydrogen, 26.7 mol % carbon monoxide, 6.4 mol % carbon dioxide and 1.7 mol % inerts (N2, Ar and CH4), based on the total number of moles in the ATR syngas.
A mixture of synthesis gas from a steam methane reforming (SMR) and a SGP synthesis gas. The mixture comprises 65.8 mol % hydrogen, 25.6 mol % carbon monoxide, 4.4 mol % carbon dioxide and 3.8 mol % inerts (N2, Ar and CH4), based on the total number of moles in the mixture of syngas.
Table 2a provides an overview of the feed, i.e. carbon dioxide ex. field, synthesis gas and hydrogen ex. combined effluent, provided to the methanol synthesis.
Table 2b provides an overview of the feed composition provided to the methanol synthesis.
Table 3 provides an overview of the feedstock, i.e. methane-comprising feed, oxygen and water, required to produce the synthesis gas.
Table 4 shows the methanol production based on field carbon dioxide.
Experiment 1a: (not According to the Invention)
The methanol feed to the OTO process is synthesised from a mixture of SGP and SMR synthesis gas. 2949 ton/day of natural gas is required to produce sufficient methanol.
Experiment 1b: (not According to the Invention)
The methanol feed to the OTO process is synthesised from a mixture of part of the hydrogen ex. combined effluent and SGP synthesis gas. 2722 ton/day of methane-comprising feed is required to produce sufficient methanol.
Furthermore, inert (N2, Ar and CH4) concentrations in the feed to the methanol synthesis are reduced, compared to the levels seen in experiment 1a, due to dilution of the SGP synthesis gas with hydrogen obtained from the ethane cracker.
Experiment 1c:
The methanol feed to the OTO process is synthesised from a mixture of carbon dioxide ex. field, hydrogen ex. combined effluent and SGP synthesis gas. By addition of field carbon dioxide, the carbon dioxide content is increased to 3.3 mol %, based on the total feed to the methanol synthesis. The methane-comprising feed consumption for producing the methanol has decreased by 12 wt % and 5 wt % respectively based on the methane-comprising feed required for producing the methanol in Experiments 1a and 1b. In addition, 255 ton/day of methanol is produced based on field carbon dioxide, i.e. not produced as part of the process to prepare synthesis gas, which would need to be sequestered or otherwise captured and stored. As a result, the carbon dioxide penalty of the process is reduced.
Again, inert (N2, Ar and CH4) concentrations are further lowered.
Experiment 1d:
The methanol feed to the OTO process is synthesised from a mixture of carbon dioxide ex. field, hydrogen ex. combined effluent, SGP synthesis gas and the addition of additional hydrogen for instance from a second or further ethane cracker unit, or styrene production unit. By addition of field carbon dioxide, the carbon dioxide content is increased to 7.9 mol %, based on the total feed to the methanol synthesis. The methane-comprising feed consumption for producing the methanol has decreased by 27 wt % and 21 wt % respectively based on the methane-comprising feed required for producing the methanol in Experiments 1a and 1b. In addition 1062 ton/day of methanol is produced based on field carbon dioxide.
Again, inert (N2, Ar and CH4) concentrations are further lowered.
Experiment 1e: (not According to the Invention)
The methanol feed to the OTO process is synthesised from a mixture of part of the hydrogen ex. combined effluent and ATR synthesis gas. 2933 ton/day of methane-comprising feed is required to produce sufficient methanol.
Experiment 1f:
The methanol feed to the OTO process is synthesised from a mixture of carbon dioxide ex. field, hydrogen ex. combined effluent and ATR synthesis gas. By addition of field carbon dioxide, the carbon dioxide content is increased to 7.1 mol %, based on the total feed to the methanol synthesis. The methane-comprising feed consumption for producing the methanol has decreased by 6 wt % and 5 wt % respectively based on the methane-comprising feed required for producing the methanol in Experiments 1a and 1e. In addition 273 ton/day of methanol is produced based on field carbon dioxide. As a result, the carbon dioxide penalty of the process is reduced.
Again, inert (N2, Ar and CH4) concentrations are further lowered.
Experiment 1g:
The methanol feed to the OTO process is synthesised from a mixture of carbon dioxide ex. field, hydrogen ex. cracker and ATR synthesis gas. By addition of field carbon dioxide from the MEG production unit the carbon dioxide content is increased to 7.9 mol %, based on the total feed to the methanol synthesis. The methane-comprising feed gas consumption for producing the methanol has decreased by 9 wt % and 8.9 wt % respectively based on the methane-comprising feed required for producing the methanol in Experiments 1a and 1e. In addition, 443 ton/day of methanol is produced based on field carbon dioxide. As a result, the carbon dioxide penalty of the process is reduced.
Again, inert (N2, Ar and CH4) concentrations are further lowered.











TABLE 1








Ethane
Integrated OTO/ethane



OTO
cracker
cracker



103 kg/day
103 kg/day
103 kg/day





















Feed:






Methanol
5012

5012



Ethane

2755
2755



Products:



Ethylene
512
2187
2713



Propylene
1275
50
1325



Ethane
17



propane
48
19
 67



>C3
318
135
 453



Fuel gas
98
197
 280



H2O
2710
−83
2627



Hydrogen
17
168
  186**







*based on CH2 in the feed



**including hydrogen obtained by recycling ethane in the effluent of the OTO to the ethane cracker.

















TABLE 2A






Feed






103 kg/day



H2 ex.


Product



combined


103 kg/day


Exp.
effluent
Synthesis gas
CO2 ex MEG
Methanol







1a
—
2236 (ex. SMR)
—
5000




3883 (ex. SGP)


1b
129
5488 (ex. SGP)
—
5000


1c
173
5208 (ex. SGP)
365
5000


1d
321
4322 (ex. SGP)
1584 
5000


1e
 82
5837 (ex. ATR)
—
5000


1f
134
5518 (ex. ATR)
398
5000


1g
165
5320 (ex. ATR)
635
5000


















TABLE 2b








Composition




mol %*














Exp.
H2
CO
CO2
N2, AR, CH4
H2O
Molar ratio
















1a
65.8
25.6
4.4
3.8
0.2
2.05


1b
66.1
29.7
1.8
2.2
0.2
2.04


1c
66.7
27.8
3.3
2.0
0.2
2.04


1d
68.5
21.8
8.0
1.6
0.1
2.03


1e
68.1
24.6
5.9
1.4
>0.1
2.04


1f
68.6
22.9
7.2
1.3
>0.1
2.04


1g
69.0
21.9
7.9
1.2
>0.1
2.05





*Based on the total number of moles in the feed


**molar ratio = (#mol H2 − #mol CO2)/(#mol CO + #mol CO2)

















TABLE 3






Feed


Total



103 kg/day


feedstock


Exp.
methane-comprising feed
O2
H2O
103 kg/day







1a
2949
2291
1210
6450


1b
2722
3239
  0*
5961


1c
2583
3074
  0*
5657


1d
2144
2551
  0*
4695


1e
2933
3109
1691
7733


1f
2773
2939
1599
7311


1g
2674
2834
1541
7049





*In principle no water is added.

















TABLE 4









Feed
Product




103 kg/day
103 kg/day












Exp.
CO2 ex field
Methanol
Methanol from field CO2
















1a
—
5000




1b
—
5000



1c
365
5000
255



1d
1584 
5000
1062



1e
—
5000



1f
398
5000
273



1g
635
5000
443










By using a synthesis gas such as SGP synthesis gas, which comprises relatively low amounts of carbon dioxide it is possible to capture significant amounts of field carbon dioxide in the form of methanol, ethylene, MEG or products derived therefrom. In addition, by using a synthesis gas which is produced by a process in which in principle no or only little water is used, such as a non-catalytic partial oxidation process, water consumption is significantly lowered. By using hydrogen ex. Combined effluent for the production of methanol, additional amounts of field carbon dioxide can be captured in the form of methanol, ethylene, MEG or products derived thereof.

## Claims

1. A process for producing olefins, comprising:
a. providing a feed comprising at least methane, ethane and carbon dioxide;
b. separating the feed into at least a methane-comprising feed, an ethane-comprising feed and a carbon dioxide-comprising feed;
c. providing at least part of the methane-comprising feed to a process for preparing synthesis gas to obtain a synthesis gas;
d. cracking the ethane-comprising feed in a cracking zone under cracking conditions to obtain a cracking zone effluent comprising at least olefins and hydrogen;
e. providing at least part of the carbon dioxide-comprising feed and at least part the synthesis gas obtained in step c) to an oxygenate synthesis zone and synthesising oxygenates;
f. converting at least part of the oxygenates obtained in step (e) in an oxygenate-to-olefin zone to obtain an oxygenate-to-olefin zone effluent comprising at least olefins and hydrogen;
g. combining at least part of the cracking zone effluent and at least part of the OTO zone effluent to obtain a combined effluent;
h. separating hydrogen from the combined effluent and providing at least part of the hydrogen to the oxygenate synthesis zone in step (e).

2. A process according to claim 1, comprising providing carbon dioxide, synthesis gas, and hydrogen to the oxygenate synthesis zone in a amount such that the molar ratio of hydrogen, carbon monoxide and carbon dioxide provided to the oxygenate synthesis zone is in the range of from 2.0 to 3.0.

3. A process according to claim 2, wherein carbon dioxide is present in a concentration in the range of from 0.1 to 25 mol %, based on the total number of moles of carbon dioxide, hydrogen, and carbon monoxide.

4. A process according to claim 1, wherein the olefins obtained in step (d) and/or (f) include ethylene, the process further comprising:
providing at least part of the ethylene to an ethylene oxidation zone together with a feed containing oxygen and performing a process for preparing ethylene oxide to obtain ethylene oxide and carbon dioxide; and
providing at least part of the obtained carbon dioxide to the oxygenate synthesis zone.

5. A process according to claim 1, wherein the olefins obtained in step (d) and/or (f) include ethylene, the process further comprising:
converting at least part of the ethylene with benzene to obtain styrene monomer and hydrogen; and
providing at least part of the obtained hydrogen to the oxygenate synthesis zone.

6. A process according to claim 1, wherein feed comprising at least methane, ethane and carbon dioxide is obtained from natural gas or associated gas.

7. A process according to claim 1, wherein feed comprising at least methane, ethane and carbon dioxide comprises in the range of from 0.1 to 70 mol % of carbon dioxide, based on the total content of the feedstock.

8. An integrated system for producing olefins, which system comprises:
a. a separation system, having at least an inlet for a feed comprising methane, ethane and carbon dioxide, and an outlet for a methane-comprising feed, an outlet for an ethane-comprising feed and an outlet for a carbon dioxide-comprising feed;
b. a partial oxidation system arranged to partially oxidise the methane-comprising feed to a synthesis gas, having an inlet for a methane-comprising feed to receive the methane-comprising feed from the separation system, an inlet for oxygen, and an outlet for synthesis gas;
c. a steam cracking system having one or more inlets for an ethane-comprising feed to receive the methane-comprising feed from the separation system and an inlet for steam, and an outlet for a cracker effluent comprising olefins;
d. an oxygenate-to-olefins conversion system, having one or more inlets for receiving an oxygenate feedstock, and comprising a reaction zone for contacting the oxygenate feedstock with an oxygenate conversion catalyst under oxygenate conversion conditions, and an outlet for an oxygenate-to-olefins effluent comprising olefins, including at least ethylene;
e. a work-up system arranged to receive at least part of the cracker effluent and at least part of the oxygenate-to-olefins effluent to obtain a combined effluent, the work-up section comprising a separation system, an outlet for ethylene and an outlet for hydrogen; and
f. an oxygenate synthesis system having one or more inlets for synthesis gas to receive the synthesis gas from the partial oxidation system, an inlet for hydrogen and an inlet for a carbon dioxide-comprising feed to receive the carbon dioxide-comprising feed from the separation system, and an outlet for an oxygenate feedstock; and means for providing the oxygenate feedstock from the outlet for oxygenate feedstock of the oxygenate synthesis system to the oxygenate feedstock inlet of the oxygenate-to-olefins conversion system and means for providing hydrogen from the outlet for hydrogen of the work-up section to the inlet for hydrogen of the oxygenate synthesis system.