Patent Application
20130319911
The invention relates to a method for converting hydrocarbon feedstocks comprising a shale oil into lighter products which can be utilized as fuels and/or raw materials for petrochemistry. The invention relates more particularly to a method for converting hydrocarbon feedstocks comprising a shale oil that comprises a step of hydroconverting the feedstock in an ebullating bed, followed by a step of fractionating by atmospheric distillation to give a light fraction, naphtha fraction and gas-oil fraction and to give a fraction heavier than the gas-oil fraction, a step of liquid/liquid extraction of the fraction heavier that the gas-oil fraction, and a dedicated hydrotreating for each of the naphtha and gas-oil fractions. This method enables shale oils to be converted into very-high-quality fuel bases, and is aimed more particularly at an excellent yield.
In view of high barrel price volatility and a reduction in discoveries of conventional petroleum fields, petroleum groups are turning towards non-conventional sources. Next to petroleum-bearing sands and deep offshore, bituminous shales, although relatively poorly known, are becoming ever more coveted.
Bituminous shales are sedimentary rocks which contain an insoluble organic substance called kerogen. By heat treatment in situ or ex situ (“retorting”) in the absence of air at temperatures of between 400 and 500° C., these shales liberate an oil, shale oil, with a general appearance like that of crude petroleum.
Although of a different composition from crude petroleum, shale oils may constitute a substitute for the latter and also a source of chemical intermediates.
Shale oils cannot be directly substituted into the applications of crude petroleum. Indeed, although these oils resemble petroleum in certain respects (for example, in a similar H/C ratio), they differ in their chemical nature and in a much greater level of metallic and/or non-metallic impurities, thereby making the converting of this non-conventional resource much more complex than that of petroleum. Shale oils have, in particular, levels of oxygen and of nitrogen that are much higher than those in petroleum. They may also contain higher concentrations of olefins, of sulphur or of metal compounds (especially compounds containing arsenic).
Shale oils obtained by pyrolysis of kerogen contain a large number of olefinic compounds resulting from cracking, and this translates into additional hydrogen demand at the refining stage. For instance, the bromine index, which enables calculation of the concentration by weight of olefinic hydrocarbons (by addition of bromine to the ethylenic double bond), is generally greater than 30 g/100 g of feedstock for shale oils, whereas it is between 1 and 5 g/100 g of feedstock for residues of petroleum. The olefinic compounds resulting from cracking are essentially composed of monoolefins and diolefins. The unsaturations present in the olefins are a potential source of instability by polymerization and/or oxidation.
The oxygen content is generally higher than in heavy crudes, and may be as much as 8% by weight of the feedstock. The oxygen compounds are often phenols or carboxylic acids. Consequently, shale oils may have a marked acidity.
The sulphur content varies between 0.1% and 6.5% by weight, necessitating stringent desulphurizing treatments in order to meet the specifications for fuel bases. The sulphur compounds are in the form of thiophenes, sulphides or disulphides. Moreover, the sulphur distribution profile within a shale oil may be different from that obtained in a conventional petroleum.
The most distinctive feature of the shale oils, nevertheless, is their high nitrogen content, which makes them unsuitable as a conventional feedstock for the refinery. Petroleum generally contains around 0.2% by weight of nitrogen, whereas crude shale oils contain generally of the order of 1% to approximately 3% by weight or more of nitrogen. Moreover, the nitrogen compounds present in petroleum are generally concentrated in relatively high boiling ranges, whereas the nitrogen of the compounds present in crude shale oils is generally distributed throughout all of the boiling ranges of the material. The nitrogen compounds in petroleum are primarily non-basic compounds, whereas, generally, around half of the nitrogen compounds present in crude shale oils are basic. These basic nitrogen compounds are particularly undesirable in refinery feedstocks, since these compounds often act as catalyst poisons. Furthermore, the stability of the products is a problem which is common to numerous products derived from shale oil. Such instability, including photosensitivity, appears to result essentially from the presence of nitrogen compounds. Consequently, crude shale oils must generally be subjected to a stringent refining treatment (high total pressure) in order to obtain a synthetic crude petroleum or fuel base products which meet the specifications in force.
It is also known that shale oils may contain numerous metal compounds in traces, generally present in the form of organometallic complexes. The metal compounds include the conventional contaminants such as nickel, vanadium, calcium, sodium, lead or iron, but also metal compounds of arsenic. Indeed, shale oils may contain an amount of arsenic of more than 20 ppm, whereas the amount of arsenic in crude petroleum is generally in the ppb (parts per billion) range. All of these metal compounds are catalyst poisons. More particularly, they irreversibly poison the hydrotreating catalysts and hydrogenating catalysts by gradually being deposited on the active surface. The conventional metal compounds and part of the arsenic are found primarily in heavy cuts, and are removed by deposition on the catalyst. On the other hand, when the products containing arsenic are capable of generating volatile compounds, these compounds may be found partly in the lighter cuts and may, as a result, poison the catalysts in subsequent converting processes, during refining or in petrochemistry.
Furthermore, shale oils generally contain sandy sediments originating from bituminous shale fields from which the shale oils are extracted. These sandy sediments may give rise to clogging problems, especially in fixed bed reactors.
Lastly, shale oils contain waxes, which give them a pour point higher than the ambient temperature, thereby preventing their transport in oil pipelines.
In view of appreciable resources, and in view of their evaluation as being a promising source of petroleum, there exists a genuine need for converting shale oils into lighter products which can be utilized as fuels and/or raw materials for petrochemistry. Methods for converting shale oils are known. Conventionally, conversion is practised alternatively by coking, by hydrovisbreaking (thermal cracking in the presence of hydrogen) or by hydroconverting (catalytic hydrogenation). Liquid/liquid extraction processes are also known.
For instance, U.S. Pat. No. 4,483,763 describes a method for converting shale oils with the aim of reducing their nitrogen content. This method includes a step of partial hydrogenation followed by a step of liquid/liquid extraction with a mixture of a polar organic solvent, an acid and water. The extraction is carried out either on a middle distillates cut (400-680° F.=204-360° C.), or on all of the discharge obtained by hydrogenation.
U.S. Pat. No. 5,059,303 describes a method for converting shale oils which comprises a step of hydroconverting in an ebullating bed or fixed bed, an optional fractionating step, a step of liquid/liquid extraction on a liquid fraction or on the entirety of the liquid discharge, with a solvent, thereby allowing the condensed aromatics to be extracted. The raffinate obtained after evaporation of the solvent is subsequently subjected to fractionation to give a middle distillates fraction containing up to 1000 ppm of nitrogen, and a heavier fraction containing from 500 to 3000 ppm of nitrogen. U.S. Pat. No. 5,059,303 also describes a variant of the method, which comprises a step of hydroconverting in an ebullating bed, a step of gas/liquid separation without pressure reduction, a step of liquid/liquid extraction of the liquid phase, and a step of hydrotreating of the gaseous phase.
The particular feature of shale oils in having a certain number of metallic and/or non-metallic impurities makes it much more complex to convert this non-conventional resource than petroleum. The challenge for the industrial development of methods for converting shale oils is therefore the need to develop methods which are suited to the feedstock, allowing the yield of high-quality fuel bases to be maximized. The conventional refining treatments known from petroleum must therefore be adapted to the specific composition of the shale oils.
The present invention aims to improve the known methods for converting hydrocarbon feedstocks comprising a shale oil by increasing, especially, the yield of fuel bases for a combination of steps having a specific linkage, and a treatment appropriate to each fraction obtained from the shale oils. Likewise, an object of the present invention is to obtain high-quality products having more particularly a low sulphur, nitrogen and arsenic content, preferably meeting the specifications. Another objective is to provide a method which is simple, i.e. having as few steps as necessary, while remaining effective, allowing capital investment costs to be limited.
In its broadest form, the present invention is defined as a method for converting hydrocarbon feedstock comprising at least one shale oil having a nitrogen content of at least 0.1%, often at least 1% and very often at least 2% by weight, characterized in that it comprises the following steps:
a) The feedstock is treated in a section for hydroconverting in the presence of hydrogen, said section comprising at least one ebullating bed reactor operating in gas and liquid upflow mode and containing at least one supported hydrotreating catalyst,
b) The effluent obtained in step a) is conveyed at least partly, and often entirely, into a fractionating zone, from which, by atmospheric distillation, a gaseous fraction, a naphtha fraction, a gas-oil fraction and a fraction heavier than the gas-oil fraction are recovered,
c) Said naphtha fraction is treated at least partly, and often entirely, in a section for hydrotreating in the presence of hydrogen, said section comprising at least one fixed bed reactor containing at least one hydrotreating catalyst,
d) Said gas-oil fraction is treated at least partly, and often entirely, in another section for hydrotreating in the presence of hydrogen, said section comprising at least one fixed bed reactor containing at least one hydrotreating catalyst,
e) The fraction heavier than the gas-oil fraction is subjected to a liquid/liquid extraction to give a raffinate and an extract.
The treating section in step a) typically comprises from one to three, and preferably two, reactors in series, and the treating section in steps c) and d) also comprises from one to three reactors in series.
The research work carried out by the applicants into the conversion of shale oils has led to the surprising finding that an improvement to the existing methods, in terms of yield of fuel bases and in terms of product purity, is possible through a combination of various steps linked in a specific way and a subsequent treatment section for each fraction obtained by the method.
The first step comprises hydroconversion in an ebullating bed. The technology of the ebullating bed, relative to the technology of the fixed bed, enables the treatment of feedstocks which are heavily contaminated with metals, heteroatoms and sediments, such as the shale oils, while exhibiting conversion rates which are generally greater than 50%. Indeed, in this first step, the shale oil is converted into molecules which enable the generation of future fuel bases. The majority of the metallic compounds, of the sediments and of the heterocyclic compounds is removed. The effluent emerging from the ebullating bed therefore contains the most resistant nitrogen and sulphur compounds, and possibly volatile arsenic compounds which are present in lighter components.
The effluent obtained in the hydroconverting step is subsequently fractionated by atmospheric distillation, producing various fractions, for which a treatment specific to each fraction is carried out subsequently. The key step in the method is that of carrying out a fractionation by atmospheric distillation before the liquid/liquid extraction step, in order to maximize separately the lighter fractions (naphtha, gas-oil), subsequently necessitating a moderate hydrotreating treatment which is adapted to each fraction, and to minimize the fraction heavier than the gas-oil fraction, necessitating a more severe treatment by liquid/liquid extraction. Thus the atmospheric distillation enables the preparation, in a single step, of the various fractions desired (naphtha, gas-oil), thereby facilitating downstream hydrotreating adapted to each fraction and, consequently, the direct production of gas-oil or naphtha fuel base products which meet the various specifications. Fractionation after hydrotreating is therefore not necessary.
Owing to the high level of reduction in contaminants in the ebullating bed, the light fractions (naphtha and gas-oil) contain fewer contaminants and can therefore be treated in a fixed bed section, which generally has improved hydrogenation kinetics in relation to the ebullating bed. Similarly, the operating conditions can be milder because of the limited contaminants content. Providing a treatment for each fraction permits better operability in accordance with the desired products. Depending on the operating conditions selected (more or less stringent), it is possible to obtain either a fraction which can be conveyed to a fuels pool or a finished product which meets the specifications (sulphur content, smoke point, cetane, aromatics content, etc.) in force.
Upstream of the catalytic hydrotreating beds, the fixed bed hydrotreating sections preferably comprise specific guard beds for any arsenic compounds and silicon compounds contained within the diesel and/or naphtha fractions. The arsenic compounds, which have escaped the ebullating bed (because they are generally relatively volatile), are trapped in the guard beds, thus preventing poisoning of the downstream catalysts, and enabling production of highly arsenic-depleted fuel bases.
The atmospheric distillation also enables the concentration of the most resistant nitrogen compounds in the fraction which is heavier than the gas-oil fraction, thereby limiting the amount to be treated by liquid/liquid extraction. The equipment and also the amount of solvent required in the liquid/liquid extraction step are thus minimized.
The fraction heavier than the diesel fraction that is obtained from the fractionating step is subjected to a liquid/liquid extraction by means of a polar solvent. The solvent used is a solvent for preferentially extracting aromatic compounds. Since the resistant residual nitrogen is located commonly in the aromatic compounds, the liquid/liquid extraction step therefore enables a reduction in the aromatic nitrogen compounds that are resistant to hydrodeazotization (deazotization by catalytic hydrogenation). It is important to stress that, in contrast to the prior art, the liquid/liquid extraction is performed solely on the heavy fraction, in order to avoid losses in yield of fuel bases during the recovery of the solvent following extraction. The products it is desired to extract from the heavy fraction preferably have a boiling point greater than the boiling point of the solvent, in order to avoid any loss of yield during the separation of the solvent from the raffinate after the extraction. The reason is that, during the separation of the solvent from the raffinate, any compound having a boiling point less than the boiling point of the solvent will unavoidably leave with the solvent and will therefore lower the amount of the raffinate obtained (and hence the yield of fuel bases). In the case of furfural as the extraction solvent, for example, having a boiling point of 162° C., the C10− compounds, compounds which are representative of the petrol/naphtha fraction, will be lost. By treating solely the heavy fraction comprising compounds having boiling points greater than the boiling point of the extraction solvent, there is no loss of these C10− compounds. Moreover, contamination of the solvent with the C10− compounds is avoided, as are the possible steps of treatment of this solvent for the purpose of its recycling. The recovery of the solvent is therefore more efficient and economical.
Another advantage of the method is the fact that the raffinate obtained from the liquid/liquid extracting step e), following evaporation of the solvent, is preferably conveyed to a catalytic cracking section [step f)], in which it is treated under conditions which enable production of a gaseous fraction, a petrol fraction, a gas-oil fraction, and a residual heavy fraction, which is referred to as “slurry”. This variant enables the yield of fuel bases to be maximized.
Another advantage is the fact that the extract obtained from the liquid/liquid extraction may be at least partly recycled to the hydroconverting step a). The recycling enables an increase in the yield of fuel bases.
The Hydrocarbon Feedstock
The hydrocarbon feedstock comprises at least one shale oil or a mixture of shale oils. The term “shale oil” is used here in its broadest sense and is intended to include any shale oil or a shale oil fraction which contains nitrogenous impurities. This includes crude shale oil, whether obtained by pyrolysis, by solvent extraction or by other means, or shale oil which has been filtered to remove the solids, or which has been treated by one or more solvents, chemical products, or other treatments, and which contains nitrogenous impurities. The term “shale oil” also comprises the shale oil fractions obtained by distillation or by another fractionating technique.
The shale oils used in the present invention generally have a Conradson carbon content of at least 0.1% by weight and generally at least 5% by weight, an asphaltenes content (IP143 standard/with C7) of at least 1%, often at least 2% by weight. Their sulphur content is generally at least 0.1%, often at least 1% and very often at least 2%, and even up to 4% or even 7% by weight. The amount of metals they contain is generally at least 5 ppm by weight, often at least 50 ppm by weight, and typically at least 100 ppm by weight or at least 200 ppm by weight. Their nitrogen content is generally at least 0.5%, often at least 1% and very often at least 2% by weight. Their arsenic content is generally greater than 1 ppm by weight, and up to 50 ppm by weight.
The method according to the present invention is intended for converting shale oils. Nevertheless, the feedstock may further comprise, in addition to the shale oil, other, synthetic liquid hydrocarbons, more particularly those which contain a substantial amount of cyclic organic nitrogen compounds. This includes oils derived from coal, oils obtained on the basis of heavy tars, bituminous sands, pyrolysis oils from ligneous residues such as wood residues, crudes obtained from biomass (“biocrudes”), vegetable oils and animal fats.
Other hydrocarbon feedstocks may also supplement the shale oil. The feedstocks are selected from the group consisting of vacuum distillates and direct distillation residues, vacuum distillates and unconverted residues obtained from conversion processes, such as, for example, those originating from distillation to the point of coke (coking), products obtained from fixed-bed hydroconversion of heavy fractions, products obtained from ebullating-bed processes for hydroconversion of heavy fractions, and oils deasphalted using solvents (for example, oils deasphalted with propane, with butane and with pentane, originating from the deasphalting of vacuum residues from direct distillation or of vacuum residues obtained from hydroconversion processes). The feedstocks may further comprise light cycle oil (LCO) of various origins, heavy cycle oil (HCO) of various origins, and also gas-oil cuts which originate from catalytic cracking and have in general a distillation range from about 150° C. to about 650° C. The feedstocks may also comprise aromatic extracts obtained in the manufacture of lubricating oils. The feedstocks may also be prepared and used in a mixture, in any proportions.
Hydrocarbons added to shale oil or to the mixture of shale oils may represent from 20% to 60% by weight of the total feedstock (shale oil or mixture of shale oils+added hydrocarbons), or from 10% to 90% by weight.
Hydroconversion
According to the present invention, the feedstock is first of all subjected to an ebullating-bed hydroconverting step [step a)]. By hydroconverting is meant reactions of hydrogenation, hydrotreating, hydrodesulphurization, hydrodenitrogenation, hydrodemetallation and hydrocracking.
The operation of the ebullating-bed catalytic reactor, including the recycling of the liquids from the reactor to the top through the agitated catalyst bed, is generally well known. Ebullating bed technologies use supported catalysts, generally in the form of extrudates having a diameter of generally of the order of 1 mm or less than 1 mm, for example greater than or equal to 0.7 mm. The catalysts remain inside the reactors and are not evacuated with the products. The catalytic activity can be held constant by virtue of on-line replacement (addition and withdrawal) of the catalyst. There is therefore no need to shut down the unit in order to change the spent catalyst, or to increase the reaction temperatures along the cycle in order to compensate for deactivation. Moreover, working with constant operating conditions enables consistent product qualities and consistent yields to be obtained throughout the cycle of the catalyst. Since the catalyst is held in agitation by substantial recycling of liquid, the head loss over the reactor remains low and constant, and the heat of reaction is rapidly averaged over the catalyst bed, which is therefore almost isothermal and does not require cooling via the injection of quenches. Implementing the hydroconversion in an ebullating bed obviates the problems of catalyst contamination that are associated with the deposits of impurities that are present naturally in shale oils.
The conditions in step a) of treating the feedstock in the presence of hydrogen are customarily conventional conditions for ebullating-bed hydroconversion of a liquid hydrocarbon fraction. It is customary to operate under a total pressure of 2 to 35 MPa, preferably of 10 to 20 MPa, at a temperature of 300° C. to 550° C. and often of 400° C. to 450° C. The hourly space velocity (HSV) and the hydrogen partial pressure are important factors, which are selected according to the characteristics of the product to be treated and to the desired conversion. The HSV is usually situated within a range from 0.2 h−1 to 1.5 h−1 and preferably from 0.3 h−1 to 1 h−1. The amount of hydrogen mixed with the feedstock is customarily from 50 to 5000 normal cubic metres (Nm3) per cubic metre (m3) of liquid feedstock, and usually from 100 to 1000 Nm3/m3, and preferably from 300 to 500 Nm3/m3.
This hydroconverting step a) may usually be implemented under the conditions of the T-STAR® process, as described for example in the article Heavy Oil Hydroprocessing, published by the AIChE, Mar. 19-23, 1995, Houston, Tex., paper number 42d. It may also be implemented under the conditions of the H-OIL® process, as described for example in the article published by NPRA, Annual Meeting, Mar. 16-18, 1997, J. J. Colyar and L. I. Wilson under the title The H-Oil®Process, A Worldwide Leader In Vacuum Residue Hydroprocessing.
The hydrogen required for the hydroconversion (and for the subsequent hydrotreating operations) may come from the steam reforming of hydrocarbons (methane) or else from the gas obtained from oil shales during the production of shale oils.
The catalyst in step a) is preferably a conventional granular hydroconversion catalyst, comprising, on an amorphous support, at least one metal or metal compound having a hydrodehydrogenating function. Generally speaking, a catalyst is used whose pore distribution is suitable for the treatment of feedstocks containing metals.
The hydrodehydrogenating function may be provided by at least one group VIII metal selected from the group consisting of nickel and/or cobalt, optionally in combination with at least one group VIB metal selected from the group consisting of molybdenum and/or tungsten. It is possible, for example, to use a catalyst comprising from 0.5% to 10% by weight of nickel and preferably from 1% to 5% by weight of nickel (expressed as nickel oxide, NiO) and from 1% to 30% by weight of molybdenum, preferably from 5% to 20% by weight of molybdenum (expressed as molybdenum oxide, MoO3), on an amorphous inorganic support. The total amount of oxides of metals from groups VIB and VIII is often from 5% to 40% by weight and generally from 7% to 30% by weight and the weight ratio expressed as metal oxide between group VI metal (or metals) and group VIII metal (or metals) is generally from 20 to 1 and usually from 10 to 2.
The support of the catalyst will be selected, for example, from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also include other compounds, for example oxides selected from the group consisting of boron oxide, zirconia, titanium oxide and phosphoric anhydride. It is usual to use an alumina support, and very often an alumina support doped with phosphorus and optionally with boron. In this case, the concentration of phosphoric anhydride, P2O5, is customarily less than about 20% by weight and usually less than about 10% by weight, and at least 0.001% by weight. The concentration of boron trioxide, B2O3, is customarily from approximately 0% to approximately 10% by weight. The alumina used is customarily a y (gamma) or η (eta) alumina. This catalyst is usually in the form of an extrudate. The catalyst in step a) is preferably based on nickel and molybdenum, doped with phosphorus and supported on alumina. Use may be made, for example, of an HTS 458 catalyst sold by Axens.
Prior to the injection of the feedstock, the catalysts used in the method according to the present invention preferably undergo a sulphurizing treatment to convert at least partly the metallic species into sulphides before they are contacted with the feedstock to be treated. This activation treatment by sulphurization is well known to the skilled person and may be carried out by any method already described in the literature, either in situ, i.e. within the reactor, or ex situ.
The spent catalyst is partly replaced with fresh catalyst by withdrawal at the bottom of the reactor and introduction at the top of the reactor of fresh or new catalyst at regular intervals, for example by individual or quasi-continuous addition. It is possible, for example, to introduce fresh catalyst every day. The level of replacement of the spent catalyst by fresh catalyst may be, for example, from approximately 0.05 kg to approximately 10 kg per m3 of feedstock. This withdrawal and this replacement are carried out using devices which allow the continuous operation of this hydroconverting step. The unit customarily comprises a recirculation pump for maintaining the catalyst in an ebullating bed by continuous recycling of at least part of the liquid withdrawn at the top of the reactor and reinjected at the bottom of the reactor. It is also possible to convey the spent catalyst withdrawn from the reactor into a regenerating zone, in which the carbon and sulphur it contains are removed, and then to return this regenerated catalyst to the hydroconverting step a).
The operating conditions coupled with the catalytic activity allow feedstock conversion rates of possibly from 50% to 95%, preferably from 70% to 95%, to be obtained. The aforementioned degree of conversion is defined as the mass fraction of the feedstock at the start of the reaction section minus the mass fraction of the heavy fraction having a boiling point of more than 343° C. at the end of the reaction section, this figure being divided by the mass fraction of the feedstock at the start of the reaction section.
The technology of the ebullating bed allows treatment of feedstocks which are highly contaminated with metals, sediments and heteroatoms, without facing head loss problems or clogging problems, which are known when a fixed bed is used. The metals, such as nickel, vanadium, iron and arsenic, are largely removed from the feedstock by deposition on the catalysts during the reaction. The remaining (volatile) arsenic will be removed in the hydrotreating steps by specific guard beds. The sediments present in the shale oils are also removed via the replacement of the catalyst in the ebullating bed without disrupting the hydroconversion reactions. These steps also enable the removal, by hydrodenitrogenation, of the major part of the nitrogen, leaving only the most resistant nitrogen compounds.
The hydroconversion in step a) enables an effluent to be obtained that contains not more than 3000 ppm, preferably not more than 2000 ppm, by weight of nitrogen.
Fractionation by Atmospheric Distillation
The effluent obtained in the hydroconverting step is conveyed at least partly, and preferably in its entirety, into a fractionating zone, from which a gaseous fraction, a naphtha fraction, a gas-oil fraction and a fraction heavier than the gas-oil fraction are recovered by atmospheric distillation.
The effluent obtained in step a) is preferably fractionated by atmospheric distillation into a gaseous fraction having a boiling point of less than 50° C., a naphtha fraction boiling at between about 50° C. and 150° C., a gas-oil fraction boiling at between about 150° C. and 370° C., and a fraction which is heavier than the gas-oil fraction and which boils generally at above 340° C., preferably at above 370° C.
The naphtha and diesel fractions are subsequently conveyed separately into hydrotreating sections. The heavy fraction undergoes a liquid/liquid extraction.
The gaseous fraction contains gases (H2, H2S, NH3, H2O, CO2, CO, C1-C4 hydrocarbons, etc.). It may advantageously undergo a purifying treatment for recovery of the hydrogen and its recycling into the hydroconverting section in step a) or into the hydrotreating sections in steps c) and d). Following purifying treatments, the C3 and C4 hydrocarbons may be used to form LPG (liquefied petroleum gas) products. The uncondensable gases (C1-C2) are generally used as internal fuel for the heating ovens of the hydroconversion and/or hydrotreating reactors.
Hydrotreating of the Naphtha Fraction and of the Gas-Oil Fraction
The naphtha and gas-oil fractions are subsequently subjected separately to fixed-bed hydrotreating [steps c) and d)]. Hydrotreating refers to reactions of hydrodesulphurization, hydrodenitrogenation and hydrodemetallation. The objective, depending on the operating conditions, which are selected so as to be more or less stringent, is to bring the various cuts up to the specifications (sulphur content, smoke point, cetane, aromatics content, etc.) or to produce a synthetic crude petroleum. Treating the naphtha fraction in one hydrotreating section and the gas-oil fraction in another hydrotreating section allows improved operability in terms of the operating conditions, so as to be able to bring each cut up to the required specifications with a maximum yield and in a single step per cut. In this way, fractionation after hydrotreating is unnecessary. The difference between the two hydrotreating sections is based more on differences in operating conditions than on the selection of the catalyst.
The fixed-bed hydrotreating sections preferably comprise, upstream of the catalytic hydrotreating beds, specific guard beds for the arsenic compounds (arsenic-containing compounds) and silicon compounds that are optionally present in the naphtha and/or diesel fractions. The arsenic-containing compounds which have escaped the ebullating bed (being generally relatively volatile) are trapped in the guard beds, thereby preventing the poisoning of downstream catalysts and enabling highly arsenic-depleted fuel bases to be obtained.
The guard beds which allow removal of arsenic and silicon from naphtha or gas-oil cuts are known to the skilled person. They comprise, for example, an absorbent material comprising nickel deposited on an appropriate support (silica, magnesia or alumina) as described in FR2617497, or else an absorbent material comprising copper on a support, as described in FR2762004. Mention may also be made of the guard beds sold by Axens: ACT 979, ACT 989, ACT 961, ACT 981.
The operating conditions in each hydrotreating section are adapted to the feedstock to be treated. The operating conditions for hydrotreating the naphtha fraction are generally gentler than those for the gas-oil fraction.
In the naphtha fraction hydrotreating step [step c)] it is customary to operate under an absolute pressure of 4 to 15 MPa, often of 10 to 13 MPa. The temperature during this step c) is customarily from 280° C. to 380° C., often from 300° C. to 350° C. This temperature is customarily adjusted in accordance with the desired level of hydrodesulphurization. The hourly space velocity (HSV) is usually situated within a range from 0.1 h−1 to 5 h−1, and preferably from 0.5 h−1 to 1 h−1. The amount of hydrogen mixed with the feedstock is customarily from 100 to 5000 normal cubic metres (Nm3) per cubic metre (m3) of liquid feedstock, and usually from 200 to 1000 Nm3/m3, and preferably from 300 to 500 Nm3/m3. It is useful to operate in the presence of hydrogen sulphide (for the sulphurizing of the catalyst), and the hydrogen sulphide partial pressure is customarily from 0.002 times to 0.1 times, and preferably from 0.005 times to 0.05 times, the total pressure.
In the gas-oil fraction hydrotreating step [step d)] it is customary to operate under an absolute pressure of 7 to 20 MPa, often of 10 to 15 MPa. The temperature during this step c) is customarily from 320° C. to 450° C., often from 340° C. to 400° C. This temperature is customarily adjusted depending on the desired level of hydrodesulphurization. The mass hourly velocity ((t of feedstock/h)/t of catalyst) is between 0.1 and 1 h−1. The hourly space velocity (HSV) is usually situated within a range from 0.2 h−1 to 1 h−1, and preferably from 0.3 h−1 to 0.8 h−1. The amount of hydrogen mixed into the feedstock is customarily from 100 to 5000 normal cubic metres (Nm3) per cubic metre (m3) of liquid feedstock, and usually from 200 to 1000 Nm3/m3, and preferably from 300 to 500 Nm3/m3. It is useful to operate in the presence of hydrogen sulphide, and the hydrogen sulphide partial pressure is customarily from 0.002 times to 0.1 times, and preferably from 0.005 times to 0.05 times, the total pressure.
In the hydrotreating sections, the ideal catalyst must have a high hydrogenating power, so as to produce thorough refining of the products, and to obtain a substantial lowering of the sulphur content and nitrogen content. In the preferred embodiment, the hydrotreating sections operate at relatively low temperature, which promotes thorough hydrogenation and a limitation on the coking of the catalyst. The use of a single catalyst or of two or more different catalysts, simultaneously or successively, in the hydrotreating sections would not depart from the scope of the present invention. The hydrotreating in steps c) and d) is customarily carried out industrially in one or more reactors with liquid downflow.
In the two hydrotreating sections [steps c) and d)], the same type of catalyst is used; the catalysts in each section may be identical or different. At least one fixed bed of conventional hydrotreating catalyst is used, comprising, on an amorphous support, at least one metal or metal compound having a hydrodehydrogenating function.
The hydrodehydrogenating function may be provided by at least one group VIII metal selected from the group consisting of nickel and/or cobalt, optionally in combination with at least one group VIB metal selected from the group consisting of molybdenum and/or tungsten. It is possible, for example, to use a catalyst comprising from 0.5% to 10% by weight of nickel and preferably from 1% to 5% by weight of nickel (expressed as nickel oxide, NiO) and from 1% to 30% by weight of molybdenum, preferably from 5% to 20% by weight of molybdenum (expressed as molybdenum oxide, MoO3), on an amorphous inorganic support. The total amount of oxides of metals from groups VI and VIII is often from about 5% to about 40% by weight, and generally from about 7% to 30% by weight, and the weight ratio expressed in terms of metal oxide between metal (or metals) from group VIB to metal (or metals) from group VIII is in general from about 20 to about 1, and usually from about 10 to about 2.
The support is for example selected from the group consisting of alumina, silica, silica-aluminas, magnesia, clays and mixtures of at least two of these minerals. This support may also include other compounds and for example oxides selected from the group consisting of boron oxide, zirconia, titanium oxide and phosphoric anhydride. It is usual to use an alumina support, and very often an alumina support doped with phosphorus and optionally with boron. In this case, the concentration of phosphoric anhydride, P2O5, is customarily less than about 20% by weight and usually less than about 10% by weight and at least 0.001% by weight. The concentration of boron trioxide, B2O3, is customarily from approximately 0% to approximately 10% by weight. The alumina used is customarily a γ (gamma) or η (eta) alumina. This catalyst is usually in the form of beads or extrudates.
Prior to the injection of the feedstock, the catalysts used in the method according to the present invention are preferably subjected to a sulphurizing treatment enabling to convert at least partly the metallic species into sulphides before they are contacted with the feedstock to be treated. This activation treatment by sulphurization is well known to the skilled person and may be carried out by any method already described in the literature, either in situ, i.e. within the reactor, or ex situ.
The hydrotreating in step c) of the naphtha cut produces a cut containing not more than 1 ppm by weight of nitrogen, preferably not more than 0.5 ppm of nitrogen, and not more than 5 ppm by weight of sulphur, preferably not more than 0.5 ppm of sulphur.
The hydrotreating in step d) of the gas-oil cut produces a cut containing not more than 100 ppm of nitrogen, preferably not more than 20 ppm of nitrogen, and not more than 50 ppm of sulphur, preferably not more than 10 ppm of sulphur.
Liquid/Liquid Extraction
The fraction heavier than the gas-oil fraction that is obtained from the atmospheric-distillation fractionating section is subsequently sent to a liquid/liquid extracting step [step e)]. The objective in this step is to extract the aromatic compounds, including the resistant nitrogen from the heavy fraction, to give a raffinate which can be used as a feedstock for the catalytic cracking in a conventional fluid-bed catalytic cracking unit. This therefore makes it possible to maximize the yield of fuel bases. Accordingly, the liquid/liquid extraction enables value to be derived from a fraction which conventionally is too resistant to be hydrotreated.
The extraction is performed by means of a solvent which is known for preferential extraction of aromatic compounds. As the solvent it is possible to use furfural, N-methyl-2-pyrrolidone (NMP), sulpholane, dimethylformamide (DMF), dimethyl sulphoxide (DMSO), phenol, or a mixture of these solvents in equal or different proportions.
The liquid/liquid extraction may be carried out by any means known to the skilled person. The extraction is generally carried out in a mixer-settler or in an extraction column. The extraction is preferably carried out in an extraction column.
The operating conditions are generally a solvent/feedstock ratio of 1/1 to 3/1, preferably of 1/1 to 1.8/1, a temperature profile of between the ambient temperature and 150° C., preferably between 50° C. and 150° C. The pressure is located between the atmospheric pressure and 2 MPa, preferably between the atmospheric pressure and 1 MPa.
The solvent selected has a boiling point which is sufficiently high to allow the heavy fraction obtained from the fractionation to be fluidified without evaporating, the heavy fraction being typically carried over at temperatures between 200° C. and 300° C.
Following contact of the solvent with the heavy fraction, two phases are formed: (i) the extract, composed of parts of the heavy fraction that are not soluble in the solvent (and with a high concentration of aromatics containing resistant nitrogen), and (ii) the raffinate, composed of the solvent and soluble parts of the heavy fraction, which constitutes a feedstock from which value can be derived by catalytic cracking in order to enhance the yield of fuel bases. The solvent is separated by distillation from the soluble parts, and is recycled internally to the liquid/liquid extraction process; the management of the solvent is known to the skilled person.
The extraction produces a raffinate containing not more than 1500 ppm, preferably not more than 1000 ppm, of nitrogen. At least a part, and preferably the entirety, of the raffinate obtained from the liquid/liquid extraction is preferably conveyed to a catalytic cracking step.
According to one preferred variant, at least part, and preferably the entirety, of the extract obtained in liquid/liquid extracting step e) is recycled to the start of step a).
According to another variant, the extract is conveyed into an oxyvapogasification section, in which it is converted into a gas containing hydrogen and carbon monoxide. This gaseous mixture can be used for the synthesis of methanol or for the synthesis of hydrocarbons by the Fischer-Tropsch reaction. This mixture, in the context of the present invention, is preferably conveyed into a “shift” conversion (steam conversion) section in which, in the presence of steam, it is converted into hydrogen and into carbon dioxide. The hydrogen obtained may be employed in steps a), c) and d) of the method according to the invention. The extract obtained in step e) may also be used as solid fuel or, after fluxing, as liquid fuel, or may form part of the composition of bitumens and/or of heavy fuel oils.
The liquid/liquid extraction of the heavy fraction therefore enables extraction of the resistant aromatic compounds containing nitrogen, and of the contaminants (metals). Carrying out the extraction solely on the heavy fraction makes it possible to prevent losses of feedstock for the catalytic cracking, and therefore to increase the overall yield of the method. Recycling the extract to hydroconverting step a) also allows the yield to be increased.
Catalytic Cracking
Finally, according to one abovementioned variant, in a catalytic cracking step [step f)], at least part, and preferably the entirety, of the raffinate, obtained in step e), may be conveyed, after solvent evaporation, into a conventional catalytic cracking section, in which said raffinate is treated conventionally, under conditions well known to the skilled person, to produce a gaseous fraction, a petrol fraction, a gas-oil fraction and a heavy fraction, referred to as “slurry”. The gas-oil fraction will for example be conveyed at least partly to fuel reservoirs (pools) and/or recycled, at least partly, or even in its entirety, to the gas-oil hydrotreating step d). The heavy fraction will, for example, be at least partly, or even in its entirety, conveyed to the heavy fuel oil reservoir (pool) and/or recycled at least partly, or even in its entirety, to the catalytic cracking step f). In the context of the present invention, the expression “conventional catalytic cracking” encompasses cracking processes which comprise at least one step of catalyst regeneration by partial combustion, and those which comprise at least one step of catalyst regeneration by total combustion, and/or those comprising both at least one partial combustion step and at least one total combustion step.
For example, a summary description of catalytic cracking (the first industrial implementation of which goes back to 1936 (Houdry process) or 1942 for the use of fluidized bed catalyst) will be found in Ullmans Encyclopedia of Industrial Chemistry Volume A 18, 1991, pages 61 to 64. It is customary to use a conventional catalyst comprising a matrix, optionally an additive and at least one zeolite. The amount of zeolite is variable but is customarily from about 3% to 60% by weight, often from about 6% to 50% by weight and usually from about 10% to 45% by weight. The zeolite is customarily dispersed in the matrix. The amount of additive is customarily from about 0% to about 30% by weight. The amount of matrix represents the rest up to 100% by weight. The additive is generally selected from the group consisting of the oxides of metals from group IIA of the periodic table of the elements, such as, for example, magnesium oxide or calcium oxide, the oxides of rare earths, and the titanates of the metals from group IIA. The matrix is usually a silica, an alumina, a silica-alumina, a silica-magnesia, a clay or a mixture of two or more of these products. The zeolite most commonly used is zeolite Y. Cracking is carried out in a substantially vertical reactor in either upflow or downflow mode. The selection of the catalyst and of the operating conditions are dependent on the target products in dependence on the feedstock treated, as is described, for example, in the article by M. Marcilly, pages 990-991, published in the journal of the Institut Français du Pétrole, November-December 1975, pages 969-1006. It is customary to operate at a temperature from 450° C. to 600° C. and with reactor residence times of less than 1 minute, often from about 0.1 to about 50 seconds.
The catalytic cracking step f) may also be a fluidized bed catalytic cracking step, for example according to the process called R2R. This step may be performed conventionally as known to skilled persons under appropriate cracking conditions for producing hydrocarbon products with a lower molecular weight. Descriptions of operation and of catalysts which can be used in the context of fluidized bed cracking in this step f) are described for example in the patent documents U.S. Pat. No. 5,286,690, U.S. Pat. No. 5,324,696 and EP-A-699224.
The fluidized bed catalytic cracking reactor may operate in upflow mode or in downflow mode. Although not a preferred embodiment of the present invention, it is likewise possible to contemplate performing the catalytic cracking in a moving bed reactor. Particularly preferred catalytic cracking catalysts are those containing at least one zeolite, customarily in a mixture with an appropriate matrix such as, for example, alumina, silica or silica-alumina.
According to
In
A shale oil is treated that has the characteristics set out in Table 1.
The shale oil is treated in an ebullating bed reactor containing the commercial catalyst HOC 458 from Axens. The operating conditions are as follows:
The liquid products obtained from the reactor are fractionated by atmospheric distillation to give a naphtha fraction (C5+-150° C.), a gas-oil fraction (150-370° C.) and a residual fraction 370° C.+.
The naphtha fraction is subjected to fixed bed hydrotreating using an NiMo-on-alumina catalyst. The operating conditions are as follows:
The gas-oil fraction is subjected to fixed bed hydrotreating using an NiMo-on-alumina catalyst. The operating conditions are as follows:
The residual fraction is subjected to a liquid/liquid extraction with furfural, with a solvent/feedstock ratio of 1.8/1, at a temperature of 100° C. and at atmospheric pressure. This gives a raffinate and an extract.
The raffinate is subsequently subjected to catalytic cracking using a catalyst containing 20% by weight of zeolite Y and 80% by weight of a silica-alumina matrix. This feedstock, preheated to 135° C., is contacted at the bottom of a vertical reactor with a catalyst from a regenerator, the catalyst having been regenerated under hot conditions. The entry temperature of the catalyst into the reactor is 720° C. The ratio of the catalyst flow rate to the feedstock flow rate is 6.0. The calorific input of the catalyst at 720° C. enables the evaporation of the feedstock and the cracking reaction, both of which are endothermic. The average residence time of the catalyst in the reaction zone is approximately 3 seconds. The operating pressure is 1.8 bar absolute. The catalyst temperature measured at the end of the upwardly driven (riser) fluidized bed reactor is 525° C. The cracked hydrocarbons and the catalyst are separated by virtue of cyclones situated in a stripping zone (stripper) in which the catalyst is stripped. The catalyst, which has become loaded with coke during the reaction and then has been stripped in the stripping zone, is subsequently conveyed into the regenerator. The coke content of the solid (delta coke) at the start of the regenerator is 0.85%. This coke is burnt by air injected into the regenerator. The combustion, which is very exothermic, raises the temperature of the solid from 525° C. to 720° C. The hot regenerated catalyst emerges from the regenerator and is conveyed back to the bottom of the reactor.
The hydrocarbons separated from the catalyst emerge from the stripping zone. They are sent to a main fractionating tower, from which the gases and petrol cuts emerge at the top, and then, at the bottom of the tower, in order of increasing boiling point, the LCO and HCO cuts and the slurry (370° C.+) emerge.
Table 2 gives the properties of the various feedstocks in each step and also the yields obtained in the various units, and the overall yield. Hence it is observed that, starting from 100% by weight of shale oil, 87.2% by weight of products (LPG, naphtha, middle distillates) are obtained conforming to the commercial Euro V specifications.
| Number | Date | Country | Kind |
|---|---|---|---|
| FR 10 61244 | Dec 2010 | FR | national |
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/FR11/53020 | 12/16/2011 | WO | 00 | 8/21/2013 |
