Patent Application
20080289999
The present invention relates to the oil industry and the use of sulphur-containing heavy refinery residues. Oil is traditionally treated in the oil refinery by a set of fractionation and chemical conversion operations to produce a set of final commercial products satisfying well defined standards or specifications, for example distillation ranges, sulphur contents, characteristic technical indices such as the octane number or diesel number, etc.
The principal final commercial products are petrochemical naphtha, gasoline, kerosene, gas oil (also termed diesel fuel), domestic fuel, and other categories of fuel with greater or lesser sulphur contents, road asphalt, liquefied petroleum gas, and sometimes other products: lubricating oils, solvents, paraffin, gas turbine fuel, etc. An oil refinery thus produces a relatively large number of final commercial products from a certain number of crude oils, selected as a function of their composition and their price.
Changes in markets, particularly increasing competition from natural gas, and specifications regarding discharges from burning facilities (discharge of oxides of sulphur, oxides of nitrogen, solid particles, in particular in Europe) has led to a sharp reduction in the markets for sulphur-containing heavy fuels, for example heavy fuel containing at most 3.5% or 4% of sulphur. Thus, refiners are confronted with a major technical problem, that of using sulphur-containing refinery residues, while satisfying regulatory requirements. Such sulphur-containing fuels are typically excessive and many states are tending to limit the sulphur, in fuels to 1% of sulphur and in future to 0.5% or even 0.3%.
A further tendency in the use of oil products is a tendency to increase the consumption of middle distillates and gasoline to the detriment of fuel, the increase in the consumption of middle distillates tending to be greater than that of gasoline.
The invention concerns a process for pre-refining oil, typically in the production region, to obtain one or more improved quality oils as final products (of pre-refining). These (pre-refined) oils are typically evacuated, sold and transferred to refineries for refining.
The Applicant's French patent application FR-04/02.088 already proposes the use of a field gas, which is typically cheap, to pre-refine a conventional oil and typically produce an oil Pa with a low sulphur content substantially free of asphaltenes, and also a residual oil Pb (comprising the starting asphaltenes, partially converted by a hydrogenating treatment). After refining, the oil Pa produced shall produce very little or no sulphur-containing fuel, and may have a high middle distillate content, market demand for which is increasing. This is a high quality oil. The oil Pb typically comprises inferior quality fractions, in particular residual asphaltenes.
The process also envisages the possible co-production of commercial final oil products: naphtha, gas oil, etc.
That prior art process thus produces a high quality oil Pa demanded by the market. However, there is still a need for further improvement to the pre-refining process, in particular to reduce the quantity of residual oil Pb and/or the sulphur content of its heavy fractions, with low reaction volumes to reduce the cost of the pre-refining unit, for market matching and for high quality oil, to satisfy market demands and to further upgrade the proposed product.
U.S. Ser. No. 10/369,869 discloses a high performance process for converting residues separately converting pure asphalt (65% conversion) and deasphalted oil to 90% conversion, to obtain an overall residue conversion of close to 80% by weight. That process describes an ebullated bed hydroconversion process typically with 2 reactors in series (see the figure and description of that figure).
Further, a pure asphalt hydroconversion process is described in FR-A-2 396 065. The conversion of products boiling below 524° C. (or vacuum distillates and lighter) is not known, and the process described uses only a single asphalt hydroconversion reactor.
The invention proposes a process for pre-refining crude oil, generally conventional, which can produce at least two pre-refined oils Pa and Pc, wherein Pa is a high quality oil, which is non-asphaltenic (substantially free of asphaltenes) which, after refining, will produce a plurality of high value products (naptha, gasoline, middle distillates). In contrast, Pc is a residual oil (having a residue containing asphaltenes), and after refining will produce large quantities of fuel.
According to an essential characteristic of the invention, ebullated bed conversion of a heavy asphaltenic stream ASF, which is typically a stream with a very high concentration of virgin asphaltenes, is carried out. ASF generally comprises at least 40% and usually at least 50% by weight of virgin asphalt (the term “virgin” meaning that the compounds derive directly from crude oil, neither cracked nor hydrotreated). In contrast to the prior art, it has surprisingly been found that deep conversion of asphalt (>60%, for example 65% or more) is certainly possible and practical, in particular when using at least 2 reactors in series, for example 2, 3 or 4 reactors in series (or more), but that this deep conversion, naturally desired by the refiner as it agrees with changes in the market, is not desirable and that marginal conversion beyond about 56% conversion provides limited advantages in terms of increments in upgradable products and the change in the quality of products (in particular sulphur, nitrogen and metals in effluents), but in contrast require additional reaction volumes and an astonishingly high hydrogen consumption.
In other words, it has been found that it is more advantageous, in a given reaction volume, to supply a much larger quantity of feed, which results in a reduction in the conversion and desulphurization (which are desirable), but with a large increase in the total quantity of distillable products produced (lower concentration but higher quantity) as well as in the total quantity of sulphur eliminated (in the form of H2S).
We have also found that it is advantageous to keep at least two and preferably at least three hydroconversion steps in series even though a priori, the target of relatively low conversion renders this option illogical.
We have also found that it is more advantageous to carry out a treatment of the deasphalted oil, typically separated, with a substantial hydrotreatment but a medium conversion, not of the order of 90%. In accordance with one preferred characteristic variation of the invention, the residual DAO with a low sulphur content (substantially hydrotreated), is then available in a sufficient quantity to both supply a makeup with a low sulphur content to the hydroconversion effluent of the ASF stream (to compensate for its relatively high sulphur content), and to constitute one of the components of the high quality non-asphaltenic oil Pa.
Finally, it has been found that good hydroconversion results are obtained along with a better hydrodynamic behaviour in the first reactor, typically not requiring particular management of the gas fluidization conditions if a suitable diluent is used to fluidize the asphalt contained in the ASF.
The invention proposes an oil pre-refining process which, from a crude oil P (or several oils P1, P2, etc), can produce at least two oils Pa and Pc, or even 3 pre-refined oils Pa, Pb, Pc. Pa and Pb are typically high quality oils, i.e. oils substantially free of asphaltenes, while Pc is a residual oil comprising asphaltenes. The process may also possibly produce more than 3 pre-refined oils, for example 3, or more, asphaltene-free oils. It may also produce 2 or 3 residual oils (containing asphaltenes), or more. The scope of the present invention does not exclude the process from co-producing final or refined products: fuel, naptha, kerosene, gas oil, domestic fuel, oils or oil bases, etc.
The crude oil(s) P is/are typically conventional, for example Arabian light, but it is also possible to use any type of crude oil, conventional, heavy, asphaltenic, and in particular any oil with an API of 5 to 50.
The process uses fractionation of oil P by at least initial distillation (termed “atmospheric”) and generally a vacuum distillation. The initial distillation preferably separates it into at least two fractions, one being relatively richer in middle distillates and the other being relatively richer in naphtha (or at least having different naphtha/middle distillate ratios).
The process also typically comprises at least one hydrotreatment unit HDT or a conversion unit, in particular vacuum distillate hydrocracking VGO.
It comprises an atmospheric residue or vacuum residue deasphalting unit and a unit for hydrotreatment and/or hydrocracking of the deasphalted oil DAO produced.
Finally, it comprises a unit for hydroconversion of a very heavy feed, ASF, comprising at least 50% by weight of virgin asphalt, and generally substantially no or no non-virgin asphalt (already hydroconverted and recycled).
In general, the invention proposes a process for pre-refining at least one crude oil P, comprising:
All of the cuts from P (typically with hydrotreatment or hydrocracking of the vacuum distillate VGO and optional hydrotreatment of middle distillates) may be combined into a single residual asphaltenic oil Pc.
Preferably, however, optionally after hydrogenating catalytic treatment(s), at least one non-asphaltenic pre-refined oil Pa can be produced from non-asphaltenic cuts from P. These cuts may be hydrocracked and/or hydroconverted and/or hydrotreated (each may optionally be treated separately in accordance with one of the possible variations for hydrogenating catalytic treatment; this term will be explained below).
ASF is a very heavy feed: the solvent used is preferably relatively heavy (in particular heavier than propane) and thus produces an asphalt which is concentrated in asphaltenes. Suitable solvents all principally comprise paraffinic hydrocarbons (optionally olefinic) containing 3 to 7 carbon atoms. However, highly preferably, they comprise propane-butane mixtures, butane, pentane, hexane, heptane, light gasoline and mixtures obtained from the solvents cited above. Preferred solvents comprise butane, pentane, hexane and mixtures thereof. Highly preferred solvents comprise butane, pentane and mixtures thereof.
The solvent SDA deasphalting operation may be operated under conventional conditions: reference should be made to the article by BILLON et al published in 1994 in volume 49, No 5 of the review by the Institut Français du Pétrole, pp 495-507, in the book “Raffinage et conversion des produits lourds du pétrole” [refining and converting heavy oil products] by J F Le Page, S G Chatila and M Davidson, TECHNIP publications, pp 17-32, or to the description given in FR-B-2 480 773 or in FR-B-2 681 871 or in U.S. Pat. No. 4,715,946. Deasphalting may in particular be carried out at a temperature in the range 60° C. to 250° C. with one of the solvents cited above, optionally supplemented with an additive. The solvents used and the additives are in particular described in the documents cited above and in the following patent documents: U.S. Pat. No. 1,948,296; U.S. Pat. No. 2,081,473; U.S. Pat. No. 2,587,643; U.S. Pat. No. 2,882,219; U.S. Pat. No. 3,278,415 and U.S. Pat. No. 3,331,394. The solvent may be recovered by vaporization or distillation, or by an opticritical process, i.e. under supercritical conditions. Deasphalting may be carried out in a mixer-decanter or in an extraction column.
As mentioned above, RHDC conversion in ebullated beds in series of ASF is carried out in a relatively moderate manner, or even to a very low level for a process with several steps and reactors in series, using a relatively high flow rate of feed and/or a relatively low reaction volume as regards the supply flow rate.
At the cost of a relatively small loss as regards targets (in particular desulphurization), this can greatly reduce the reaction volume (and thus cost) as well as the hydrogen consumption.
The pressure used for hydroconversion of the ASF stream is typically in the range 12 to 26 MPa, in particular in the range 14 to 21 MPa. The temperature is typically in the range 340° C. to 460° C., usually in the range 390° C. to 450° C. It should be noted that, rather than mild operating conditions, relatively severe conditions are used, typically at a maximum temperature which can control the reactions and avoid too many secondary reactions (condensation) as well as too rapid deactivation of the catalyst by coking. The degree of conversion per unit volume of reactor is thus relatively high, and the final conversion is only moderate because the reaction volume is relatively low as regards the flow rate of the feed ASF. The rate by weight at which the feed is introduced per reaction stage (or per reactor) is typically at least 0.54 h−1 or even at least 0.7 h−1 and usually at least 1.08 h−1 or even at least 1.4 h−1 (expressed in kg/h of ASF feed supplied to the reactor under consideration per kg of catalyst in the reactor). This mass rate is selected to be sufficiently high so that conversion is moderate and as desired. It is typically less than 5 h−1, for example in the range 0.7 h−1 to 5 h−1, in particular in the range 1.08 h−1 to 4 h−1, in particular in the range 4 h−1 to 3 h−1.
The unexpected behaviour of the ASF feed which is very rich in virgin asphaltenes compared with hydroconversion at high conversion levels could be explained in a number of ways, without being bound to any particular one of these interpretations:
Since the feed is highly concentrated in heavy products, the unwanted heavy polynuclear aromatic compound condensation reactions are thus only slightly inhibited. This results in a negative effect on yield and product quality (sulphur, H/C ratio, nitrogen), and an increased tendency to deactivate the catalyst by coking.
It appears that the use of a concentration of virgin asphalt in the hydroconversion feed, while it allows the use of low volume reactors for a given quantity of asphaltenes (and crude oil), produces unexpected effects as regards hydroconversion.
However, said effects are largely neutralized according to the invention. The typical conversion of ASF feeds into compounds boiling below 524° C. (compared with initial products boiling at least 524° C.) is thus limited in the invention, generally to at most 58% by weight, typically to at most 56% by weight (for example in the range 30% to 56% by weight) and generally at most 54% by weight (for example between 32% and 54% by weight). Preferred conversions are at most 52% by weight, or even at most 50% by weight or even at most 48% by weight or at most 46% by weight. The minimum typical conversion is about 30% by weight, or 32% by weight or 34% by weight. In particular, a conversion in the range 34% to 52% by weight, in particular in the range 36% to 50% by weight, may be used. The minimum convention of 30% by weight should not be considered to be a threshold where a favourable technical effect occurs, but to be a limit which is to be considered minimal in order to carry out hydroconversion. Below that limit, the feed is considered to be insufficiently transformed. Further, the flow rate of the feed cannot be increased indefinitely (which would lead to conversions below 30%) as fractionation costs would become too great.
Thus, the feed flow rate (or reactor volume) is adjusted so that the conversion is in the desired range, taking the operating conditions into account. For a given reaction volume, an increase in the feed flow rate reduces conversion.
Highly preferably, and when a low conversion is desirable, the ASF is hydroconverted in a residue hydroconversion unit RHDC comprising at least 3 hydroconversion steps in series in at least 3 reactors in an ebullated bed in series, for example 3 or 4 reactors in series. It has been found that for ASF feeds highly charged with asphaltenes of the invention, the use of 2, 3 or 4 (or more) steps and reactors in series can effectively combat the loss of performance as regards desulphurization resulting from the use of a low conversion. The mean conversion per reactor of compounds boiling below 524° C. is typically less than 28% by weight, often 22% or 19% or even 17% by weight. The mean conversion per reactor of compounds boiling below 340° C. (gas oil and lighter compounds) is typically less than 15% by weight often 12% or 10% or even 8% by weight.
In a preferred variation of the invention, E2 is hydrotreated and/or hydrocracked and/or hydroconverted (HDC) at a pressure of at least 8 MPa, with a conversion of products boiling below 524° C. of at most 80% by weight, to obtain an effluent comprising a substantial fraction of hydrotreated deasphalted oil HDAO with a sulphur content reduced by at least 50% with respect to the amount of sulphur in the HAS fraction (the ASF hydroconversion effluent) boiling above 524° C., a stream comprising a first fraction F1 of said hydrotreated deasphalted oil HDAO is mixed with at least the major portion of HAS to produce a (pre-refined) asphaltenic oil Pc and a second fraction F2 of said hydrotreated deasphalted oil HDAO, optionally after supplemental hydrotreatment, is mixed with other non-asphaltenic oils derived from P, optionally hydrotreated and/or hydroconverted and/or hydrocracked to obtain a pre-refined non-asphaltenic oil Pa.
The conversion of E2 is generally at most 75%, preferably at most 70% or at most 60% or at most 50% by weight, to increase the available quantity of available residual deasphalted oil HDAO, and generally at least 40% (for example in the range 45% to 70% by weight). It may in general be carried out in an ebullated bed, typically at a pressure which is substantially lower than for ASF hydroconversion, and generally between 8 and 13 MPa. It is also possible to carry out hydroconversion or hydrotreatment in a fixed bed at a similar pressure, or even to carry out a sequence of fixed bed hydrotreatments HDT then ebullated bed hydroconversions HDC, or the reverse sequence, HDC+HDT.
The rate by weight of introducing the feed per reaction stage (or per reactor) is typically at least 0.54 h−1 or even at least 0.7 h−1 and usually at least 1.08 h−1 or even at least 1.4 h−1 (expressed in kg/h of feed E2, or optionally supplemented E2 supplied to the reactor under consideration per kg of catalyst in the reactor). This rate is selected so as to be sufficiently high for conversion to be moderate and as desired. It is typically less than 12 h−1, for example in the range 0.7 h−1 to 10 h−1, in particular in the range 1.08 h−1 to 6 h−1, in particular in the range 1.4 h−1 to 5 h−1.
In a preferred variation of the invention, the heavy asphaltenic stream ASF is not composed of pure asphalt but also comprises other lighter compounds. Typically, ASF comprises at most 95% by weight or at most 90% by weight of virgin asphalt AS or total asphalt in the case of a recycle. Advantageously, ASF comprises at least 10% or at least 12%, or even at least 15% by weight (and generally less than 50%, or even less than 40% or 35% by weight) of compounds boiling below 524° C. (typically virgin or simply hydrotreated), for example by the presence or addition of a diluent DIL essentially boiling below 565° C., or even below 524° C. Preferably, in addition to virgin asphalt AS, ASF comprises at least one diluent DIL which is liquid at ambient temperature, and comprise at least 30% by weight of compounds from P boiling below 340° C.
It has been discovered that this presence or addition of a relatively light diluent fluid has a favourable effect on the function of the reactor, in particular on the first reactor, and more so when the number of reactors in series is high (3 or more). The function of the first reactor becomes more stable, more similar to that of the subsequent reactors, and does not result in specific fluidization (increased gas flow rate to obtain an intensity of turbulence and suitable ebullated bed homogeneity). A more stable and homogeneous function improves the yields. With a diluent DIL as defined above or below, it is even possible to increase the conversion and operate at up to 60% or even 62% conversion. More moderate conversions remain advantageous, however.
The DIL compounds are usually virgin or simply hydrotreated. More generally, they are typically derived from P, either virgin or post hydrotreatment and/or hydrocracking, or after hydroconversion of DAO and/or VGO with or without prior or post hydrotreatment. The only optional treatments are thus hydrogenating catalytic treatments from the group constituted by hydrotreatments HDT, hydroconversions HDC of feeds which are typically non-asphaltenic, hydrocracking HDK (mild, medium or high pressure) and combinations of those treatments.
The use of such fractions for the diluent may be surprising: the light fractions are very low in naphthenoaromatics or polyaromatic precursors of naphthenoaromatics as required by theory (such as those which may result in the reaction medium from the use of LCO/HCO (light/heavy cycle oil) catalytic cracking fractions. Thus, it may be considered that the beneficial effect observed results principally from the effect of reduction of the viscosity of the liquid in the reaction medium.
Typically, the diluent DIL (for example essentially comprising fractions boiling below 565° C. (distillable) or 524° C.) represent between 3% and 40% by weight with respect to the asphalt included in ASF. Generally, DIL comprises between 3% and 30% by weight with respect to the asphalt in ASF, of fractions from P which are virgin or simply hydrotreated boiling below 340° C.
In a variation, DIL comprises a fraction of a composition which is substantially identical to the oil P. DIL may in particular comprise, in an amount of 4% to 40% by weight with respect to the asphalt included in ASF, a stream essentially constituted by desalted oil P. This is advantageous as it limits the quantity of oil to be fractionated, said non fractionated oil being usefully converted.
DIL may also comprise a fraction of topped desalted oil (for example P free of C5 (or C6) and lighter fractions or the fraction of P boiling above 150° C.) in the same proportions as above. It is also possible to use a mixture of middle distillates principally boiling below 340° C. and a fraction of P as cited above.
In accordance with the invention, Pa, Pc (and optionally Pb) are oils, final products from the pre-refining process each intended to be used as an initial distillation feed for one or typically more oil refineries:
Pa, Pc (or Pb) are final oils from the pre-refining process. However, these are conventional oil refinery feeds and not final products or intermediate refining products, or final products intended for a particular straight run distillation. They each typically comprise at least 6% by weight of naphtha N, at least 10% by weight of middle distillates MD (for example at least 4% by weight of kerosene [170° C./250° C.] and at least 6% by weight of diesel cut [250° C./360° C.]) and at least 10% by weight of vacuum distillate VGO.
In general, at least the major part of Pa and Pb is transported by pipelines and oil tankers for their use as initial distillation feeds for one or typically several oil refineries,
Reference will now be made to
A crude oil P, typically conventional (for example light Arabian), is supplied via a line 1 to a desalter 2. The desalted oil supplies, via a line 3, a preliminary distillation column PRE-DIST, reference 4 (often termed initial distillation or atmospheric distillation) typically functioning at a pressure in the range 0.1 to 0.5 MPa. This column, which may optionally carry out summary fractionation, produces a light stream, typically naphtha and lighter compounds, via a line 30, a stream of middle distillates MD, typically kerosene and diesel cut via line 5, and a stream of intermediate gas oil IGO via a line 6, may comprise fractions principally boiling between 340° C. to 420° C. Said intermediate gas oil, which is relatively heavy for an atmospheric column, may be obtained by major steam stripping.
The column 4 also produces an atmospheric residue via a line 7, which supplies a vacuum distillation column VAC-DIST with reference numeral 8. This column, which typically functions at a pressure in the range 0.004 to 0.04 MPa, produces a stream of vacuum distillate VGO via a line 10, and a stream of vacuum residue VR via a line 9. It may also optionally produce a stream of light vacuum distillate LVGO via a line 11.
The vacuum residue VR is supplied to a solvent deasphalting unit SDA with reference numeral 12 (preferably pentane) to produce a deasphalted oil DAO moving in the line 13 and an asphalt stream AS evacuated via a line 14.
The asphalt AS is mixed with a diluent stream DIL supplied via a line 15. This stream typically comprises a stream of desalted oil supplied via line 3 via line 15 and a stream of middle distillates MD supplied from line 5 via a line 22 and optionally a stream of intermediate gas oil IGO supplied via the line 6 via lines 23 and 22. DIL may also comprise naphtha N removed from a line 30.
Typically, ASF comprises 20% by weight of diluent DIL with respect to the virgin asphalt AS separated from deasphalting SDA and is constituted by a mixture of AS and DIL. A typical composition of DIL, compared with AS, is as follows: 10% by weight of desalted crude oil P+10% by weight of middle distillates boiling between 170° C. and 340° C. (or 360° C.).
The ASF mixture of virgin asphalt AS and diluent (fluxed asphalt) then supplies the ebullated bed hydroconversion unit RHDC with reference numeral 16. This unit typically comprises at least 2, and preferably at least 3 ebullated bed reactors arranged in series.
At the outlet from the RHDC unit, the hydroconversion effluent is supplemented by several streams moving in lines 30c, 31c, 32c, 33c and 34c. These streams typically comprise naphtha N (line 30c), hydrotreated middle distillates MD (line 31c), hydrotreated or hydrocracked (generally partially) intermediate gas oil IGO (line 32c), hydrotreated or hydrocracked (generally partially) vacuum distillate VGO (line 34c). Thus, a (pre-refined) oil Pc is reconstituted from the hydroconversion effluent, which comprises unconverted asphaltenic fractions, and typically hydrotreated or hydrocracked non-asphaltenic fractions, and thus with a reduced sulphur content. Said oil Pc has a much lower sulphur content than that of the initial oil P.
Fractions MD, IGO, VGO, DAO are then hydrotreated and/or hydrocracked (typically partially) in units H1 with reference numeral 21, H2 with reference numeral 20, H3 with reference numeral 19 and H4 with reference numeral 18. Typically, H1 (and often H2) is a hydrotreatment HDT, and H3 and H4 are mild hydrocracking units: M-HDK, medium pressure: MP-HDK, or high pressure: HP-HDK. Preferably, H4 is an ebullated bed hydrocracking unit.
The light stream moving in line 30 is subdivided into 3 elementary streams 30a, 30b, 30c.
The effluent from H1 moving in line 31 is subdivided into 3 elementary streams 31a, 31b, 31c.
The effluent from H2 moving in line 32 is subdivided into 3 elementary streams 32a, 32b, 32c.
The effluent from H3 moving in line 33 is subdivided into 3 elementary streams 33a, 33b, 33c.
The effluent from H4 moving in line 34 is subdivided into 3 elementary streams 34a, 34b, 34c.
From streams 30a, 31a, 32a, 33a and 34a, a pre-refined oil Pa is formed by mixing. Pa is an oil which is substantially free of asphaltenes since each of its components is also free of them (asphaltenes are only contained in the stream AS). It is also an oil with a very low sulphur content since the majority of its components are desulphurized, and naphtha, supplied via line 30a, is typically low in sulphur (as an option, it may also be hydrotreated).
Analogously, a pre-refined oil Pb is formed by mixing streams 30b, 31b, 32b, 33b and 34b. For the reasons given for Pa, Pb is also an oil which is substantially free of asphaltenes and has a very low sulphur content.
Advantageously, the conversions of the units are determined and the distribution of components Pa and Pb are determined so that Pa is relatively rich in gasoline and naptha precursors, and relatively low in middle distillates: kerosene and a diesel cut while Pb, in contrast, is an oil which is relatively low in gasoline and naphtha precursors, and relatively richer in middle distillates.
Thus, before re-composing oils Pa, Pb and Pc, the invention can use one or more catalytic steps using certain processes which are well known in the art, in particular desulphurizing treatments, under a pressure of hydrogen, which consume large quantities or raised quantities of hydrogen.
According to the invention, the term “hydrogenating catalytic treatment” is used to define a treatment comprising at least one of the treatments defined below symbolized by the following terms: HUT, HDC, HDK (which covers M-HDK, MP-HDK and HP-HDK), RHDT, RHDC. A hydrogenating catalytic treatment may thus comprise several of these treatments, for example HDT+HDC or HDC+HDT, etc.
The following hydrogenating catalytic treatments can thus be distinguished:
Hydrotreatment of hydrocarbon distillate or deasphalted oil (feeds substantially free of asphaltenes) are processes which are well known in the art. Their principal aim is at least partial elimination of unwanted compounds, typically sulphur, nitrogen, possibly metals such as iron, nickel or vanadium, etc. They are also often used for aromatic hydrogenation, generally simultaneously with feed desulphurization.
Conventionally, regarding those of the feeds cited above which comprise compounds boiling above 371° C., the term “hydrotreatment” defines a process wherein the conversion of these compounds into compounds with a boiling point of less than 371° C. is 20% by weight or less. For processes treating the same feeds, but with a conversion of more than 20% by weight, the term “hydroconversion” (denoted HDC), or “hydrocracking” (denoted HDK), said processes being presented below, is used.
Hydrotreatment processes functioning under hydrogen pressure use supported solid catalysts, typically granular solids or extrudates with a characteristic dimension (diameter for beads or equivalent diameter (corresponding to the same section) for extrudates) in the range 0.4 to 5 mm, in particular 1 to 3 mm. The operating conditions, in particular the space velocity (HSV) and the molar ratio of hydrogen to hydrocarbon (H2/HC), varies between cuts, the impurities present and the desired final specifications.
Non-limiting examples of the operating conditions are given in the following table:
The hydrotreatment catalysts typically comprise a metal or a compound of a metal from group VIB and a metal or compound of a metal from group VIII, on a support. The most usual catalysts are composed of an oxide support and an active phase in the form of a molybdenum or tungsten sulphide promoted by cobalt or nickel. The usual formulae used are CoMo, NiMo and NiW associations for the active phase, and γ alumina with a large specific surface area for the support. The amounts of metals are usually of the order of 9% to 15% by weight of molybdenum and 2.5% to 5% by weight of cobalt or nickel.
Certain of these catalytic formulae are sometimes doped with phosphorus. Other oxide supports are employed, such as mixed oxides of the silica-alumina or titanium-alumina type.
Said supports are typically of low acidity, to obtain acceptable catalytic cycle times.
Examples of types of catalysts and hydrotreatments, in particular diesel cuts, gas oil or vacuum gas oil cuts are catalysts HR448 and HR426 from French company AXENS.
When traces of metals, in particular nickel and vanadium, are present in the feed, a catalytic support with a porosity adapted to deposition of said metals is advantageously used.
One example of such a catalyst is HMC841 from AXENS.
For the hydrotreatment of a deasphalted oil (DAO) comprising metals, it is possible, for example, to use a first bed with a HMC841 catalyst, for demetallization, then a second bed of HR448 for desulphurization and denitrogenation.
Other technical elements relating to hydrotreatments may be found in the reference text: “Conversion processes”, by P Leprince, Technip, publishers, Paris 15th district, pages 533-574.
Hydrocracking processes are also processes which are well known in the art. They apply exclusively to feeds which are substantially free of asphaltenes or metals such as nickel or vanadium.
The hydrocracking feed is typically composed of vacuum gas oil, occasionally supplemented with gas oil and/or deasphalted oil (deasphalted vacuum residue, typically deasphalted by a solvent from the group formed by propane, butane, pentane and mixtures thereof, preferably propane and butane).
It is also possible to carry out hydrocracking of the deasphalted oil DAO. The DAO must therefore be of sufficient quality: typically, a hydrocracking feed comprises less than 400 ppm (parts per millions by weight) of asphaltenes, preferably less than 200 ppm and highly preferably less than 100 ppm. The metals contents (typically nickel+vanadium) in a hydrocracking feed are typically less than 10 ppm, preferably less than 5 ppm, and highly preferably less than 3 ppm.
Conventionally, it is considered that a feed is substantially asphaltenes-free if its asphaltenes content is less than 400 ppm. (In a similar manner, a pre-refined oil is considered to be without asphaltenes, or non-asphaltenic, if the fraction boiling above 524° C. contains less than 400 ppm of asphaltenes).
Typically, the hydrocracking feed is initially pre-refined on a hydrotreatment catalyst, which is typically different from the hydrocracking catalyst. This catalyst, typically with a lower acidity than the hydrocracking catalyst, is selected to substantially eliminate metals, reduce traces of asphaltenes, and reduce organic nitrogen, which inhibits hydrocracking reactions, to a value which is typically less than 100 ppm, preferably less than 50 ppm and highly preferably less than 20 ppm.
The hydrocracking catalysts are typically bifunctional catalysts having a double function: firstly, acidic, and secondly, hydrogenating/dehydrogenating.
Typically, the support has a relatively high acidity so that the ratio of the hydrogenating activity to the isomerizing activity, H/A, as defined in French patent FR-A-2 805 276 pages 1 line 24 to page 3 line 5, is more than 8, or preferably more than 10 or highly preferably more than 12, or even more than 15. Typically, hydrotreatment is carried out upstream of the reactor or the hydrocracking zone with a hydrotreatment catalyst wherein said ratio H/A is less than 8, in particular less than 7.
The hydrocracking catalysts typically comprise at least one metal or compound of a metal from group VIB (such as Mo, W) and a metal or compound of a metal from group VIII (such as Ni, etc) deposited on a support. The atomic ratio of metal from group VIII (MVIII) to the sum of metals from groups VIII and VIB, i.e. the atomic ratio MVIII/(MVIII+MVIB), in particular for NiMo and NiW pairs, is usually close to 0.25, for example in the range 0.22 to 0.28.
The metals content is usually in the range 10% to 30% by weight.
The group VIII metal may also be a noble metal such as palladium or platinum, in amounts of the order of 0.5% to 1% by weight.
The acidic support may comprise an alumina doped with a halogen, or a silica-alumina having a sufficiently acidity, or a zeolite, for example a dealuminated Y or USY zeolite, often having a double pore distribution with a double pore network in particular comprising micropores with a dimension principally in the range 4 to 10 Å and mesopores with a dimension principally in the range 60 to 500 Å. The silica/alumina ratio of the zeolite structure is normally in the range 6.5 to 12.
As an example, it is possible to use a sequence of hydrotreatment then hydrocracking with HR448 (HDT) then HYC642 (HDK) catalysts sold by French company AXENS. If the feed comprises metals, upstream of said two catalytic beds it is possible to use a demetallization catalyst such as the catalyst HMC841 also sold by AXENS.
Typical examples of the operating conditions for hydrocracking are as follows:
By definition, the conversion is that of products with a boiling point of more than 371° C. into products boiling below 371° C.
Typically, depending on the feeds, the partial pressure of hydrogen is usually in the range from about 2 MPa to 6 MPa for mild hydrocracking, between about 5 MPa and 10 MPa for medium pressure hydrocracking and between about 9 MPa and 17 MPa for high pressure hydrocracking. The total pressure is usually in the range 2.6 to 8 MPa for mild hydrocracking, between about 7 and 12 MPa for medium pressure hydrocracking and between 12 and 20 MPa for high pressure hydrocracking.
Hydrocracking processes are typically operated in a fixed bed with granular solids or extrudates with a characteristic dimension (diameter for beads or equivalent diameter (corresponding to the same section) for extrudates) in the range 0.4 to 5 mm, in particular in the range 1 to 3 mm. The scope of the invention encompasses hydrocracking being carried out in a moving bed (granular bed of catalyst typically in the form of extrudates or more preferably as beads), with dimensions similar to those described for a fixed bed.
Other technical elements relating to hydrocracking may be found in the reference work “Hydrocracking Science and Technology”, J Scherzer and A J Gruia, Marcel Dekker, publisher, New York, and in the reference work “Conversion processes”, P Leprince, Technip, Paris 15th district, pages 334-364.
Processes are known which can achieve conversions (with the same definition as for hydrocracking) of more than 20% by weight and frequently much higher (for example 20% to 50% or 50% to 85% by weight), for example ebullated bed processes. Said processes may use variable partial pressures of hydrogen, for example between 4 and 12 MPa, temperatures of between 380° C. and 450° C., and a hydrogen recycle which is, for example, in the range 300 to 1000 Nm3 per m3 of feed.
The catalysts used are similar or of a type close to that of catalysts for hydrotreatment or residue hydroconversion, defined below, and have a porosity which can provide a considerable demetallization capacity.
As an example, it is possible to use a catalyst of the HTS358 type, sold by French company AXENS.
Processes for hydrotreatment of residues (and hydroconversion of residues) are processes which are well known in the art.
Typical operating conditions for said processes are: hourly space velocity (HSV) in the range 0.1 to 0.5. Partial pressure of H2 between 1 and 1.7 MPa. Hydrogen recycle between 600 and 1600 Nm3 per m3 of feed. Temperature between 340° C. and 450° C.
Catalysts for fixed, moving or ebullated bed processes are usually supported macroscopic solids, for example beads or extrudates with middle distillates in the range 0.4 to 5 millimetres. Typically, they are supported catalysts comprising a metal or a metallic compound from group VIB (Cr, Mo, W), and a metal or metallic compound from group VIII (Fe, Co, Ni, etc) on a mineral support, for example catalysts based on cobalt and molybdenum on alumina, or nickel and molybdenum on alumina.
For fixed bed hydrotreatment or hydroconversion, it is possible, for example, to use a hydrodemetallization catalyst HMC841, then hydroconversion and hydrocracking catalysts: HT 318, then HT328 sold by French company AXENS.
For an ebullated bed, it is possible to use a HOC458 type catalyst, also sold by French company AXENS.
The catalysts for slurry processes are more diversified and may comprise particles of ground lignite or charcoal impregnated with iron sulphate or other metals, ground used hydrotreatment catalyst, particles of molybdenum sulphide associated with a hydrocarbonaceous matrix, obtained by in situ decomposition of precursors such as molybdenum napthenate, etc. The typical dimensions of the particles are less than 100 micrometres, or even much smaller.
Other characteristics of processes and catalysts for hydroconversion of residues are given in general reference work A: “Raffinage et conversion des produits lourds du petrole” [Refining and converting heavy products from oil] by J F Le Page, S G Chatila, M Davidson, Technip, Paris, 1990, in Chapter 4 (Conversion catalytique sous pression d'hydrogène) [catalytic conversion in hydrogen], and Chapter 3, paragraph 3.2.3. Reference could also be made to general reference work B: “Conversion processes”, P Leprince, Technip, Paris 15th District, pages 411-450, in Chapter 13 (hydroconversion de residus) [Residue hydroconversion], and to the general work: “Upgrading petroleum residues and heavy oils”, by Murray R Gray, Marcel Dekker Inc, New York, Chapter 5.
Purified gas may be used for the production of hydrogen for carrying out said hydrogenating catalytic treatments, for example by steam reforming on a nickel catalyst then steam conversion of CO followed by purification—this is a well known process, described in the work with reference B cited above, p, 451-502, or in the reference work “The desulphurization of heavy oils and residues”, J Speight, Marcel Dekker, Inc, New York.
| Number | Date | Country | Kind |
|---|---|---|---|
| 0504300 | Apr 2005 | FR | national |
| Filing Document | Filing Date | Country | Kind | 371c Date |
|---|---|---|---|---|
| PCT/FR2006/000672 | 3/24/2006 | WO | 00 | 6/11/2008 |
