PROCESSES FOR PREPARING HYDROCARBON COMPOSITIONS

Abstract

Processes, catalysts and systems for preparing a composition comprising aliphatic, olefinic, cyclic and/or aromatic hydrocarbons of seven or greater carbon atoms per molecule are provided.

Claims

1. A process for preparing a composition comprising aliphatic, olefinic, cyclic and/or aromatic hydrocarbons of five or greater carbon atoms per molecule, said process comprising: (a) providing a first hydrocarbon mixture comprising isomers of hydrocarbons; (b) providing a heterogeneous catalyst, the heterogeneous catalyst comprising: (i) pentasil zeolite in a range between about 1% to about 99%; (ii) amorphous silica, amorphous alumina, or a combination thereof, in a range of about 1% to about 99%; (iii) Zn and/or Cu, in a range between about 0.05% to about 3% by weight; and (iv) at least one exchanged metal of Group VII series in a range of about 0.05% to about 6% by weight; (c) contacting the first hydrocarbon mixture with the heterogeneous catalyst; and (d) forming a second hydrocarbon mixture comprising molecules having five or more carbon atoms, the second hydrocarbon mixture having a higher octane unit and/or a lower vapor pressure as compared to the first hydrocarbon mixture.

2. The process of claim 1, wherein the second hydrocarbon mixture has a higher octane unit of about 10 to about 25 units.

3. (canceled)

4. The process of claim 1, further comprising: stopping forming the second hydrocarbon mixture; isolating the heterogeneous catalyst; and regenerating the heterogeneous catalyst to remove carbonaceous deposits.

5. The process of claim 4, wherein the heterogeneous catalyst is at a temperature of about 200 C. to 700 C. during regenerating and the regenerating comprises passing heated wet or dry nitrogen, air or combinations thereof over the heterogeneous catalyst.

6. The process of claim 1, wherein the heterogeneous catalyst is within one or more bed members.

7. The process of claim 6, wherein each bed member comprises one or more first hydrocarbon mixture feed points.

8. The process of claim 7, wherein heated wet or dry nitrogen, air or combinations thereof is supplied to each bed catalyst bed member in parallel via the one or more first hydrocarbon mixture feed points.

9. The process of claim 7, wherein contacting the first hydrocarbon mixture with the heterogeneous catalyst further comprises using a temperature control member to control the temperature of the first hydrocarbon mixture, wherein the temperature control member is located on or about the one or more first hydrocarbon mixture feed points.

10. The process of claim 1, further comprising: transferring heat using one or more cross-exchangers or heating the first hydrocarbon mixture using one or more cross-exchangers.

11. The process of claim 1, further comprising: using a sulfur removal member comprising an adsorbent; and operating the sulfur removal member at a temperature capable of sulfur removal.

12. (canceled)

13. The process of claim 1, further comprising: obtaining the first hydrocarbon mixture from straight run naphtha derived from a crude oil distillation unit or from a natural gasoline or condensate.

14. The process of claim 1, wherein the pentasil zeolite comprises crystalline aluminosilicate silica/alumina at a mole ratio of between about 15 and about 100.

15. (canceled)

16. The process of claim 1, wherein at least one of the Group VII series metals exchanged to produce the heterogeneous catalyst is Rhenium.

17. The process of claim 10, further comprising: obtaining a liquid, semi-liquid, gaseous or semi-gaseous mixture, wherein the liquid, semi-liquid, gaseous or semi-gaseous mixture is isolated downstream of the at least one cross-exchanger.

18. The process of claim 16, further comprising isolating the second hydrocarbon mixture in a drum, absorption tower, or distillation column, or combinations thereof.

19. The process of claim 1, further comprising: contacting the first hydrocarbon mixture with the heterogeneous catalyst in step (c) at a temperature of between about 200 C. and 500 C.

20. The process of claim 1, wherein the heterogeneous catalyst is sized to achieve a weight hourly space velocity in the range about 0.01 to about 100.

21. The process of claim 1, further comprising: providing a dilution stream of liquid petroleum gas (LPG) composition comprising three and/or four carbon atoms per molecule.

22. A process for preparing a composition comprising aliphatic, olefinic, cyclic and/or aromatic hydrocarbons of five or greater carbon atoms per molecule, said process comprising: (a) providing a first hydrocarbon mixture comprising isomers of C5 or C6 hydrocarbons or combinations thereof, and, optionally, a liquid petroleum gas (LPG) composition comprising C3 or C4 hydrocarbons, or combinations thereof; (b) providing a heterogeneous catalyst, the heterogeneous catalyst comprising: (i) pentasil zeolite in a range between about 1% to about 99%; (ii) amorphous silica, alumina, or a combination thereof in a range of about 1% to about 99%; (iii) Zn and/or Cu, in a range between about 0.05% to about 3% by weight; and (iv) at least one exchanged metal of Group VII series in a range of about 0.05% to about 6% by weight; (c) contacting the first hydrocarbon mixture with the heterogeneous catalyst; and (d) forming a second hydrocarbon mixture comprising molecules having five or more carbon atoms, the second hydrocarbon mixture having a higher linear octane number and a lower vapor pressure as compared to the first hydrocarbon mixture.

23-25. (canceled)

26. The process of claim 22, wherein the pentasil zeolite comprises crystalline aluminosilicate silica/alumina at a mole ratio of between about 15 and about 100.

27-31. (canceled)

Description

BRIEF DESCRIPTION OF THE FIGURES

[0019] FIG. 1 is a process diagram depicting a nonlimiting embodiment of conversion of naphtha using a Multibed downflow reactor in accordance with the present invention.

DETAILED DESCRIPTION

[0020] This disclosure relates to processes for preparing a composition comprising aliphatic, olefinic, cyclic and/or aromatic hydrocarbons of five or greater carbon atoms per molecule. In one nonlimiting embodiment, the processes are for preparing a composition comprising aliphatic, olefinic, cyclic and/or aromatic hydrocarbons of six or greater carbon atoms per molecule. In one nonlimiting embodiment, the processes are for preparing a composition comprising aliphatic, olefinic, cyclic and/or aromatic hydrocarbons of seven or greater carbon atoms per molecule. The processes of the present invention employ a robust catalyst which can handle small quantities of sulfur, an economical choice of a fixed bed reactor design which preferably includes at least two beds of catalyst, and in some embodiments, integrated sulfur removal by adsorbent to meet latest tier 3 sulfur specifications in gasoline.

[0021] The processes of the present invention comprise providing a first hydrocarbon mixture comprising isomers of hydrocarbons. In one nonlimiting embodiment, the first hydrocarbon mixture comprises isomers of hydrocarbons containing five to seven carbon atoms. In some nonlimiting embodiments, the first hydrocarbon mixture is obtained from straight run naphtha derived from a crude oil distillation unit or from a natural gasoline or condensate.

[0022] Processes of the present invention further comprise providing a heterogeneous catalyst.

[0023] The heterogeneous catalyst comprises pentasil zeolite. In one nonlimiting embodiment, the pentasil zeolite comprises crystalline aluminosilicate silica/alumina at a mole ratio of between about 15 and about 100. In one nonlimiting embodiment, the pentasil zeolite comprises ZSM5.

[0024] In one nonlimiting embodiment, the pentasil zeolite is included in the catalyst in a range between about 1% to about 99% by weight of the total catalyst. In one nonlimiting embodiment, the pentasil zeolite is included in the catalyst in a range of between about 20% to about 85% by weight of the total catalyst. The heterogeneous catalyst further comprises amorphous silica, amorphous alumina, or a combination thereof. In one nonlimiting embodiment, the zeolite comprises a mixture, of silica and alumina.

[0025] In one nonlimiting embodiment, the silica, alumina, or a combination thereof is included in the catalyst in a range between about 1% to about 99% by weight of the total catalyst.

[0026] In addition, the heterogeneous catalyst comprises Zn, CU or both. In one nonlimiting embodiment, the Zn, Cu or both are included in a range between about 0.05% to about 3% by weight of the total catalyst.

[0027] Further, the heterogeneous catalyst comprises at least one exchanged metal of the Group VII series. In one nonlimiting embodiment, at least one of the Group VII series metals exchanged to produce the heterogeneous catalyst is Rhenium.

[0028] In one nonlimiting embodiment, the metal is in the range of about 0.05% to about 6% by weight of the total catalyst.

[0029] In one nonlimiting embodiment, the metal is in the range of about 0.5% to about 6% by weight.

[0030] In one nonlimiting embodiment, the heterogeneous catalyst is sized to achieve a weight hourly space velocity in the range about 0.01/hour to about 100/hour, where weight hourly space velocity is defined as mass flow of feed in kg/h divided by the mass of catalyst in kg.

[0031] In one nonlimiting embodiment, the first hydrocarbon mixture is contacted with the heterogeneous catalyst at a temperature of between about 200 C. and 500 C., and at a pressure of between 1 barA and 20 barA. Typically the first hydrocarbon mixture is present as a gaseous stream when it is contacted with the catalyst.

[0032] In one nonlimiting embodiment, the catalyst used within the process comprises the zeolite ZSM-5 at concentrations between 20-85% weight, Zn, Cu or both Zn and Cu at a concentration less than 3% weight and one element of Group VII at concentration less than 5% weight, along with an amorphous binder comprising silica and/or alumina. This exemplary catalyst provides high conversion of pentane or hexane at conditions between 200-400 C. and at operational pressures below 30 barg.

[0033] In one nonlimiting embodiment, the processes further comprise providing a dilution stream of liquid petroleum gas (LPG) composition comprising three and/or four carbon atoms per molecule with the first hydrocarbon mixture.

[0034] The processes further comprise contacting the first hydrocarbon mixture with the heterogeneous catalyst to form a second hydrocarbon mixture. In one nonlimiting embodiment, the second hydrocarbon mixture has a higher linear octane number. In one nonlimiting embodiment, the second hydrocarbon mixture comprises hydrocarbons having five or more carbon atoms. In one nonlimiting embodiment, the second hydrocarbon mixture comprises hydrocarbons having six or more carbon atoms. In one nonlimiting embodiment, the second hydrocarbon mixture comprises hydrocarbons having seven or more carbon atoms.

[0035] As will be understood by the skilled artisan upon reading this disclosure, in addition to higher linear octane number, motor octane units and/or research octane units can be routinely obtained.

[0036] In one nonlimiting embodiment, the second hydrocarbon mixture has a higher linear octane number or unit and a lower vapor pressure as compared to the first hydrocarbon mixture. In one nonlimiting embodiment, the second hydrocarbon mixture has a higher octane unit of about 10 units. In one nonlimiting embodiment, the second hydrocarbon mixture has a higher octane unit of about 15 units. In one nonlimiting embodiment, the second hydrocarbon mixture has a higher octane unit of about 20 units. In one nonlimiting embodiment, the second hydrocarbon mixture has a higher octane unit of about 25 units.

[0037] In one nonlimiting embodiment, the heterogeneous catalyst is within one or more bed members. Preferred, as an economical choice, is use of a fixed bed reactor design which includes at least two beds of catalyst.

[0038] In one nonlimiting embodiment, each bed member comprises one or more first hydrocarbon mixture feed points.

[0039] In one nonlimiting embodiment, the first hydrocarbon mixture is contacted with the heterogeneous catalyst using a temperature control member to control the temperature of the first hydrocarbon mixture. In one nonlimiting embodiment, the temperature control member is located on or about the one or more first hydrocarbon mixture feed points.

[0040] In some nonlimiting embodiments, the processes further comprise stopping formation of the second hydrocarbon mixture by isolating the feed on first hydrocarbon mixture from the heterogeneous catalyst. This is typically accomplished by means of one or more isolation valves in the feed-line to the catalyst. Hydrocarbons that are present in the catalyst are removed using nitrogen and or steam prior to commencing the regenerating process. The heterogeneous catalyst is regenerated to remove carbonaceous deposits. In one nonlimiting embodiment, the temperature of the catalyst during regeneration is about 200 C. to 700 C. In one nonlimiting embodiment, the regenerating comprises passing heated wet or dry nitrogen, air or combinations thereof over the heterogeneous catalyst. Typically the catalyst is regenerated by passing a stream of heated gas through the catalyst for a period of between 1 hour and 10 days. Initially the gas stream is heated nitrogen. Air is added to the nitrogen to provide an oxidant. The oxidant oxidizes carbonaceous materials which accumulate on the surface of the catalyst and affect its catalytic activity. The typical flowrate of nitrogen is 0.5-2 kg/h for every kg of catalyst. The flowrate of air is adjusted to maintain an oxygen content in the feedstream of between 1 and 20% v/v. Steam may be added to the regenerating gas stream to improve the regenerating process. The temperature of the regenerating gas stream that is supplied to the catalyst is progressively increased during regeneration process within the range 200-700 C.

[0041] In one nonlimiting embodiment, the heterogeneous catalyst is within one or more bed members. Preferred, as an economical choice, is use of a fixed bed reactor design which includes at least two beds of catalyst. In this nonlimiting embodiment, the heated wet or dry nitrogen, air or combinations thereof is supplied to each bed catalyst bed member in parallel via the one or more first hydrocarbon mixture feed points. This nonlimiting embodiment with at least two beds of catalyst can be particularly useful, whereby one is operating, while the other reactor is regenerated. These reactors are designed, such that either one may be taken off-line from the processing of the naphtha and regenerated periodically using air and nitrogen.

[0042] In some nonlimiting embodiments, the processes further comprise transferring heat using one or more cross-exchangers or heating the first hydrocarbon mixture to a temperature of between 250 and 500 C. using one or more cross-exchangers.

[0043] In some nonlimiting embodiments, the processes further comprise obtaining a liquid, semi-liquid, gaseous or semi-gaseous mixture, isolated downstream of the at least one cross-exchanger. The hydrocarbon liquid, typically comprising straight-chain, olefinic, cyclic and/or aromatic hydrocarbons of five, six, seven or greater carbon atoms per molecule, is typically obtained by either condensing the volatile components that are present in the second hydrocarbon mixture, where such condensation occurs when the stream is cooled in a heat exchange or combination of heat exchangers. The liquids may also be obtained by contacting the second hydrocarbon mixture with and absorbent liquid such as a heavier hydrocarbon stream, or a organic solvent stream. Contacting the second hydrocarbon stream with a absorbent liquid may be also be used in conjunction with a series of heat exchange units.

[0044] In some nonlimiting embodiments, the processes further comprise using a sulfur removal member, which is typically located downstream of the heterogeneous catalyst, such that it removes sulfur from the second hydrocarbon mixture. In one nonlimiting embodiment, the sulfur removal member comprises an adsorbent. In one nonlimiting embodiment, the adsorbent comprises metal oxide. In one nonlimiting embodiment, the sulfur removal member is operated at a temperature capable of removing sulfur. In one nonlimiting embodiment, the sulfur removal member is operated at a temperature of between about 80 C. to 200 C.

[0045] In some nonlimiting embodiments, the second hydrocarbon mixture is isolated in a drum, absorption tower, or distillation column, or combinations thereof.

[0046] In one nonlimiting embodiment, the present invention provides a process solution to upgrade light naphtha feedstocks, such as straight run naphtha, or condensate, and convert them into higher boiling range naphtha with increased octane number or unit and lower vapor pressure, which can be useful as gasoline blend stock.

[0047] A process diagram for the conversion of light naphtha, to form a heavier higher boiling-range material, containing aromatics, is provided in FIG. 1. The purpose of the process is to substantially convert light paraffinic compounds such as hexane and pentane, which have a high proportion of straight chain isomers, into larger paraffinic naphthenic, and aromatic components, such that they can used as high octane gasoline.

[0048] As shown in FIG. 1, the light naphtha feed comprising, for example, pentane and hexane enters the process as a liquid in stream 101. The remaining stream 102 is further split into two streams 102 and 105. Stream 102 is sent through exchangers and heaters prior to reaction. Cross-exchanger 23, uses the hot reactor effluent to heat the cold inlet feed stream, stream 102. Typical temperatures of approximately 150 to 300 C. are achieved using the cross-exchanger, resulting in stream 103. Stream 103 is further heated to a reaction temperature of approximately 200 to 500 C. using fired heater 21. Hot gas feed 104 enters the top of reactor 22 and flows downward where it is at least partially reacted over a 1.sup.st catalyst bed containing a zeolite catalyst. Cooler feed 105 may be injected into the reactor to reduce the temperature prior to being introduced into the second catalyst bed. The combined effluent from the 1.sup.st catalyst bed and cooler feed injection 105, if used, are then further reacted over a 2.sup.nd catalyst bed. Hot reactor effluent 106 exits the reactor and is cooled using cross-exchanger 23. The cooled stream 107 is processed in an optional sulfur removal bed 24, which adsorbs sulfur components. The desulfurized stream 108 is cooled in cooler 25. The resulting product stream 109, containing 2-phase liquid and vapor products are separated in vessel 26. The majority of methane and ethane, which are produced as by-products in reactor 22, exit the process in stream 110. The remaining stream is further processed in fractionation column 27 where LPG components such as propane and butane are removed as the top product 113, while the upgraded naphtha in stream 114 is directed to the gasoline pool.

[0049] In one nonlimiting embodiment, the present invention relates to a process according to FIG. 1 for preparing a composition comprising straight-chain, olefinic, cyclic and/or aromatic hydrocarbons of five, six, seven or greater carbon atoms per molecule.

[0050] An alternative nonlimiting embodiment of the naphtha upgrading process involves co-injection of LPG comprising mainly butane or propane in stream 115. An advantage of this process is that any olefin compounds present in the LPG are converted into gasoline compounds. In addition, the light olefinic material helps to initiate the upgrading reactions, increasing the octane content of the resulting naphtha.

[0051] The process flow diagram provided herein is a nonlimiting illustration of the general process. Upon reading this disclosure, the skilled artisan would understand that certain derivations known to those skilled in the art, such as, but in no way limited to, further integration with conventional FCC light-ends recovery equipment, various heat integration options and product stabilization process can be used.

[0052] The following nonlimiting examples are provided to further illustrate the present invention. A nonlimiting example of a catalyst useful in conjunction with the illustrated process and comprising pentasil zeolite in a range between about 1% to about 99%; amorphous silica, amorphous alumina, or a combination thereof, in a range of about 1% to about 99%; Zn and/or Cu, in a range between about 0.05% to about 3% by weight; and at least one exchanged metal of Group VII series in a range of about 0.05% to about 6% by weight is set forth in Example 1. Process performance of this catalyst, is disclosed by Examples 2, 3, 4 and 5.

EXAMPLES

Example 1: Catalyst

[0053] 75 grams ZSM5 powder with a Si/Al ratio of 15/1 was mixed with 20 grams of kaolin, 5 grams of carboxymethyl cellulose and enough water to form a suitable paste for extrusion after mixing in a high shear sigma blade mixer. After extrusion through a short L/D ratio multiple 3 mm trilobe geometry die plate extruder, the green extrudate was dried at 120 C. in air for 3 hours followed by calcination in air at 650-750 C. for 3 hours. After calcination, sodium exchange was performed with aqueous salts of zinc, rhenium and copper to achieve a metal loading of 1 weight percent for each metal respectively. After ion exchange, the extrudate was dried in air at 120 C. for 3 hours followed by calcination to 650-750 C. for 3 hours. After final calcination, the catalyst was prepared for activity test in a fixed bed reactor.

Example 2: 12 Hour Average Catalyst Performance, from a Time on Stream of 96-140 Hours

[0054] 10 grams of 16 mesh particle size catalyst was loaded into a 0.500 inch diameter 316 SS reactor tube, equipped with a thermocouple, located in the middle of the catalyst bed, with glass beads above and below. The reactor tube was then placed in an electric tube furnace. The reactor tube was heated to 300 C. under a constant flow of research grade nitrogen, while maintaining a back pressure of 9 barg.

[0055] Once the internal catalyst bed temperature stabilized at 300 C., the nitrogen feed was discontinued and 0.4 weight hour space velocity of n-hexane was introduced to the reactor, while maintaining a backpressure of 9 barg. The catalyst bed temperature of 300 C. was maintained for 24 hours under the constant hydrocarbon feed; after which, it was increased to 350 C. at a rate of 1 C./minute. The gaseous product stream was analyzed on-line by an Agilent 7890B gas chromatograph, while the liquid product stream was analyzed periodically offline by an Agilent 7890B gas chromatograph to confirm C4+ yield and feed conversion in the condensate. Detailed hydrocarbon analysis was obtained by employing the ASTM D6730 method. The 12 hour average catalyst performance, from a time on stream of 96-140 hours, is provided in the table below.

TABLE-US-00001 Hexane Yield to Hexane Hexane Hexane Catalyst Bed Hexane Methane/ Yield to C.sub.3 Yield to C.sub.4 Yield to C.sub.5+ Temperature WHSV Conversion ethane hydrocarbon hydrocarbon hydrocarbon 350 C. 0.4 98.2% 1.5% 15.8% 9.7% 69.0

[0056] The C.sub.5+ Product Composition was measured on a mass percentage. This GC analysis was as follows:

TABLE-US-00002 C8 Benzene Toluene aromatic 2.6% 11.9% 15.5

Example 3: 10 Hour Average Catalyst Performance from a Time on Stream of 100-110 Hours

[0057] The procedure of Example 2 was repeated for the catalyst of Example 1. The reactor was heated to 300 C., and the nitrogen feed was discontinued and 0.4 WHSV of an n-pentane was introduced at a back pressure of 9 barg. The reactor was slowly heated to 353 C. and the conditions were held constant for 100 hours. The table below is the 10 hour average catalyst performance from a time on stream of 100-110 hours.

TABLE-US-00003 Pentane Yield to Pentane Pentane Pentane Catalyst Bed n-pentane methane/ Yield to C.sub.3 Yield to C.sub.4 Yield to C.sub.5+ Temperature WHSV conversion ethane hydrocarbon hydrocarbon hydrocarbon 353 C. 0.4 96.5% 1.1% 12.5% 11.8 74.6

[0058] The C.sub.5+ Product Composition on a mass percentage basis is as follows:

TABLE-US-00004 C8 benzene toluene aromatic 3.6% 9.2% 20.2%

Example 4: Mixed Naphtha Feed Together with LPG Diluent

[0059] A reactor operated with a mixed naphtha feed together with LPG diluent. The reactor was operated at a temperature in the range 280-420 C., preferably 350 C. and a pressure of 9 barg. The feed rate was equivalent to a space velocity. Average reactor performance at the above conditions was as follows:

TABLE-US-00005 Yield to Yield to Catalyst Bed naphtha methane/ C.sub.3/C.sub.4 Yield to C.sub.5+ Octane Temperature WHSV conversion ethane hydrocarbon hydrocarbon increase 353 C. 0.4 96.5% 1.1% 15-20% 75-80% 20-25 points

Example 5: Mixed Naphtha Feed Together with LPG Diluents and a Sulfur Removal Bed

[0060] A reactor is operated with a mixed naphtha feed of 10,000-12,000 kg/h containing 90 ppm sulfur, together with LPG diluent flow of 2000 kg/h and a sulfur removal bed comprising a metal oxide. The reactor is operated at a temperature in the range 280-420 C., preferably 320 C. and a pressure of 10 barg. The sulfur removal bed is operated at a temperature of 160-180 C. The catalyst load was equivalent to a space velocity of 0.8-1.0/h. Average reactor performance at the above conditions was as follows:

TABLE-US-00006 naphtha Product Product Catalyst Bed conversion Feed Octane Sulfur Temperature WHSV to liquids octane increase content 320 C. 0.8 85% 58-60 22-25 3-5 points ppm

Example 6: 10 Hour Average Catalyst Performance from a Time on Stream of 520-530 Hours

[0061] The procedure of Example 2 was repeated for the catalyst of Example 1. The reactor was heated to 300 C., and the nitrogen feed was discontinued and 0.4 WHSV of a mixed alkane feed was introduced at a back pressure of 0.34 barg. The reactor was slowly heated to 353 C. The reactor was operated at a temperature in the range 300-350 C. for 520 hours. The reactor temperature was adjusted to 315 C. and the conditions were held constant. The table below is the 10 hour average catalyst performance from a time on stream of 520-530 hours. Feed composition:

n-hexane: 39% w/w
2-methyl pentane: 24.6% w/w
3-methyl pentane: 14.9% w/w
2,2-dimethyl butane: 2.2% w/w
2,3 dimethyl butane: 3.9% w/w

Methyl-cyclopentane: 9.9% w/w

Cyclohexane: 5.5% w/w

[0062]

TABLE-US-00007 Conversion Conversion Conversion Conversion Catalyst Bed to methane/ to C.sub.3 to C.sub.4 to C.sub.5+ Temperature WHSV ethane hydrocarbon hydrocarbon hydrocarbon 325 C. 0.4 0.4% 7.6% 6.8% 85.2%

[0063] The aromatics composition of the liquid (C5+) product on a mass percentage basis is as follows:

TABLE-US-00008 C9 benzene Toluene Xylenes aromatics 1.6% 5.7% 11.9% 8.3%