200413514 玖、發明說明： 【發明所屬之技術領域】 t明領域 本發明關於一種贺；生τ 在w石蠟烴的方法，該石蠟烴可用為 汽油摻混成分。所得釗 + 于^的n*油摻混成分係實質上不含硫化 合物。該方法包括同睥脸& 务 ^ 寺將烯烴氫化及藉由氫解將硫化合物 降解。 【先前技術】 逸今拮術维 已知道在先前技藝中有幾 及用於異丁烯的二聚合反應。 來製造異辛烷的方法係揭示於 號中。依照該方法，烴原料係 性觸媒接觸，較佳為與陽離子 部分的異丁烯係被二聚合成二 分離及氫化成異辛烧。 個方法係用於異辛烷的製造 一種由含有異丁烯的烴原料 FI 106955 和 US 6, 613, 108 於含氮化合物的存在下與酸 性交換樹脂接觸，俾至少一 異丁烯，此二異丁烯然後被 GB 1 044771揭示一種由烴原料去除硫化合物的方法， 其係藉由用-種氫化用氣體在加氫脫硫觸媒如載於氧化紹 ^上的鉑觸媒之存在下將原料中所存在的含硫化合物加氯， 該烴原料及/或氫化用氣體含有烯烴及/或氧化碳。該方法 包括使蒸氣狀的烴混合物和氫化用氣體通過加氫脫硫觸媒 ，以將所存在的硫轉化成硫化氫，及去除硫化氫。 由丁烯二聚合反應所得到的產物典型上係含有當作雜 200413514 質的微量含硫化合物。此 ^ 4^ . ^ ^ 二雜貝會影響產物異辛烷當作汽 油摻混成分的利用性。先 田 時將、膝γ > 引技藝刊物完全沒有教導一種同 打將烯烴氫化及去除含硫抑 於ψ > $ Β 的間皁步驟，而可知道對 於”亥方法有明顯之需要。 導#用#、a Θ ^外，先刖技藝刊物完全沒有教 等使用低過罝的氫/烯烴谁 含护化人物 枓比來同時將烯烴氫化及去除 【發明内容】 本發明一目的在於提供一 ^ u 種由含有由丁烯之二聚合戶;1 獲传的烯烴進料混合物來製 &八+ +、 I仏不3 &的石蠟烴之汽油摻混 双< 刀之方法。 化人：：:"明的方法包括同時將烯烴氫化及藉由氫解將硫 化“勿降解。依本發明的方法需要低過量的氫。 申：專利範圍中揭示依本發明的方法之特徵，其用於 I每不含硫的石蠟烴之汽油摻混成分。 【實施方式】 羊細說明 ▽人駕異地’已發現可達成上述目的及避免先前技藝 之方法的缺點，或在特定的操作條件下藉由本發明的方法 來實貝減少它。依本發明，使用含有8〇_97重量%的。烯 烴、"Ο重量%的。12烯烴、。卜7重量_C9、^u 較重的>(^2稀煙以及視情況地微量之 原料。原料可為起源由丁稀之二聚合反應所；得：= 200413514 合物。 原料3有丨―1 000wt —PPm的硫化合物，典型上i—5〇wt- ppm ’此係以硫、硫化物、二硫化物、噻吩及/或烷基瞳吩 來計算。 適合於本發明之方法所使用的氫化觸媒係在氧化鋁載 體上的貴金屬觸媒。較佳的貴金屬觸媒係鉑及/或鈀，尤 佳為鉑。載體上貴金屬的含量係< 1重量％。在某些案例中 ，貝金屬可被鎳所替換。原料中高辛烷值三甲基戊烯和其 匕分枝烯烴的異構化係較沒有價值的，當使用氧化鋁當作 觸媒的載體時係能抑制較少的分枝異構物。 在該方法中’原料係在兩步驟中被氫化。在第一步驟 中，大部分的C8烯烴被轉化，但較重的烯烴和硫化合物之 轉化係頗少的。在第二步驟中，其餘的Cs烯烴、烯烴 和其它較重的烯烴與硫化合物反應。硫化合物係被轉化成 輕烴及硫化氫。 在第一步驟中，產物流視需要可被循環，以便稀釋已 濃縮的烯烴進料，而因此安全地由雙鍵的飽和去除反應熱 〇 第一步驟的反應溫度係在150 — 230 °c的範圍内，而壓 力係在20-70巴的範圍内。 在第二步驟中，於反應器内施予比第一步驟中較高的 反應溫度。该反應溫度係在18 〇 - 3 0 0 °C的範圍内，成型上 在1 90-260°C的範圍内，且壓力在20-70巴的範圍内。 在第一步驟中，h烯烴被氫化，轉化率為大於7〇0/〇。 200413514 藉由使用上述揭示之比較低的反應溫度，則能避免或至少 減少在第一步驟中觸媒的失去活性。 在第一步驟中，使用固定床式三階段氫化反應器，且 車交4圭為滴流床式反應器。反應器典型上當作滴流床式反應 器來實施’當與流體床比較下，該滴流床式反應器具有超 過兩倍的流體循環流量。此控制反應器觸媒床中的溫度上 升°在滴流床式反應器中，觸媒係置於固定床中，而液態 稀進料和氣態氫進料係同向下流經過床。此三相流動的 較佳流動方式係滴流及脈動流動方式。而且，三相氫化反 應的μ動方式如逆向流動和同向上流亦可行的。 用流體循環及藉由將新鮮的氫導入第一步驟，以維持 第步驟中南的氫濃度。循環的流體係被溶解的氫所飽和 。第一步驟中高的氫濃度之優點為有利於烯烴及將硫化物 破壞，且同時有效地減少觸媒上之焦炭形成。 依本發明，在第一步驟的製程中，可使用 或數個：相平仃的反應器，或可將反應器分割成超過一個 的反應，具有冷卻器在各反應器床間。藉由使用間隔 h _ρ盗可減少循％流量。藉由熱交換器可以從製程中 t循環的束財料反額，m將反應_於預熱將 進入丁烯二聚合單元的原料，或用當作丁婦二聚合單元之 的蒸館塔之能源，為直接或間接經由工廠或現場 楚一止 ”、、了直接用於其它現場的製程單元。 一 乂驟係在另一反應器中 烯烴及較重的烯烴係被轉化 …、餘部分的c8 双馬石％fe，且同時藉由氫解 200413514 反應將其餘的硫化合物降解。在第二步驟中，反應產物並 不循環，俾保持進料儘可能濃縮。在第二步驟中，反應器 亦典型上當作滴流床式反應器來操作，但是可應用固定床 三相氫化的其它流動方式。在第二步驟中，反應容器視需 要可結合於相同於第一反應步驟的反應容器。 在反應器之後的穩定塔中，主要由 產物中分離出輕烴、硫化氫及含硫雜質 第1和2圖中更詳細說明本發明。這些圖式僅說明本 發明的較佳具體態樣，其非限制本發明的範圍。 在第1圖中，於第一步驟中，於進入初步反應器30之 前，烴進料10與新鮮的氫進料20合併。所合併的進料1〇 和20係通經初步反應器3〇中的觸媒床33,而氣流6〇與 液流31係在反應器的下部互相分開。液流31在熱交換器 80中被冷卻，該熱交換器8G用於冷卻來自丁烯之二聚^ 反應所得料程序8卜在冷卻後，將㈣32 = 環流62和流6卜該循環流62係回到合併的進料^和別 ，該流係通往第二步驟，進人第二反應器内 入第二步驟的第二反應器4。之前，流以係與 第進 步驟的氣態產物之氣流6〇合併。 弟 應器4◦的下部互相分開，且它們係通往穩二= 氫和=成::與石蟻產物流52分開而成為流心 在第2圖中，揭不本發明的另一具 步驟中的初步反應器俜被& ^ Λ 羨，其中第一 刀割成二個觸媒床，具有液體產 10 200413514 物用出口介於觸媒床之間。在 係被冷卻及回到第二觸媒床。 器中。 第一觸媒床後，液態產物流 視情況地，冷卻可設於反應 在第2圖中，於第一牛 — 弟v驟中’在初步反應器30之前， 煙進料1 0與新鮮的氫進料 α合併。合併的進料10和20 係通經反應為30中的第一總^甘― 、 τ刃弟觸媒床34。在觸媒床34後，氣 流與液流分開’而液供Q 1 、χ 、" 、、 ” K主用於冷卻的熱交換器9 0。 冷卻後的液流9 2係回到第-網上甘 、 j弟一觸媒床35，然後氣流60與液 流31在反應器的下部互相分 1立不目刀開。熱交換器9〇可如同熱交 換器80地操作，使用來自 、 用术自丁烯之二聚合的程序流當作冷 卻流體。以來丁稀之二哿人斛 一 1 口所传到的程序流81，使液流31 在熱交換器8 0中被洽;§ρ。妒你 溉令部然後，在冷卻後，液流32被分 割成為循環流6 2和流β〗，你土、2 才机61，後者通往第二步驟進入第二反 應裔4： 0内。於進人笛一丰_ 士 c $入第一步驟中的第二反應器40之前，流 1係與含有來自第一步驟的氣態產物之氣流6〇合併。合 =抓6M 61然後通往反應器4()中的觸媒床以，而氣 7 0和液流71係在；5虛哭/1 η λα 在反應益40的下部互相分開，且它們通 往穩定塔50，1中备釦甘〜4一上、、 - 飞#八匕輕成为與石蠟產物流52分開 成為流51。 當與現今技術水準比較下，依本發明用於製造汽油摻 混成分的方法係具有數個優點。硫化合物的去除及稀煙飽 和作用係發生在於溫和操作條件τ的同—反應器内。使用 :有少於1重量%的在氧化紹載體上之貴金屬於氫化 重的c12烯烴兩者。使用氧化鋁载體於貴金屬觸媒上會 11 200413514 防止分枝的高辛烷值汽油成分異構化成較少分枝的低辛燒 值成分，其在使用氧化鋁/矽石載體時係會可觀地發生。 第一步驟中反應器的較低反應溫度會減少觸媒的失活作用 。再者，第一步驟中反應器中的產物流之循環，以及起源 於任何氫化反應器的中間束流之循環，確保溫度的安全控 制。在第二步驟中，沒有循環，而濃縮的進料確保其餘未 反應的物質之轉化。 本發明又一優點為可以使用與所必須之化學計算消耗 量接近的氫/烯烴比例，以便達成所需要的烯烴轉化及去 除硫。氫/烯烴莫耳比係〇·9至2 〇，較佳1〇5至15，此 導致低的氫消耗。低過量的氫及溫和的操作條件會減少氫 隨蒸氣產物的損失，而且通常不需要使用氫循環壓縮機於 氫回收。 以下實施例更進一步說明本發明，然而其非視為本發 明範圍的限制。 實施例 實施例1 在兩步驟方法中烯烴之氫化及硫化合物之去除，及所得到 的產物之品質，當作汽油摻混成分 在進料中以硫計算的13wt — ppm硫化合物之q —烯烴 進料，係在使用鉑觸媒（具有〇·3重量％鈀在氧化鋁載體上） 兩步驟方法中被氫化。於兩步驟中進行氫化，其中在第一 ν驟中施予6倍於新鮮進料流的循環流，而且第二步驟係 不需要任何循環流來操作。在第一步驟中，溫度係約180 12 200413514 C，而在第二步驟係220°C。 在第一步驟中，辛烷的轉化率係92%，且2, 4, 4-三甲 基戍稀對異率燒的選擇率係98%。進料含有5重量％的十二 烯。30重量％的2,2,6,6一四—4一亞甲基-庚烷被轉化， 2’ 2, 4, 6, 6-五甲基-庚烯（反式）及2, 2, 4, 6, 6一五甲基—3一庚 稀（順式）係未完全轉化。 在第二步驟中，於220°C辛烯的轉化為約1〇〇%，十二 烯的轉化率為83-94%。總烯烴轉化率於是為04%，且 Ο Α Λ ，，4—三甲基戊烯對異辛烷的選擇率係100%。硫含量係 '、的Ww卜PPm減低至產物中的0· 6wt-ppm。使用氣相 ;iff 士、 y〜丨 、9一 法分析最終產物的成分。分析結果（P〇NA=石蠟、烯 、 裒垸、芳煙）係在以下表1中，而其它結果係在表2 中0200413514 (1) Description of the invention: [Technical field to which the invention belongs] The present invention relates to a method for generating paraffin hydrocarbons, which can be used as gasoline blending ingredients. The resulting n * oil blending components of Zhao + Yu were substantially free of sulfur compounds. The method includes hydrogenation of olefins and degradation of sulfur compounds by hydrogenolysis. [Prior art] It has been known in the past that there are several techniques for dimerization of isobutene in the prior art. The method for producing isooctane is disclosed in No. According to this method, it is preferable that the hydrocarbon raw material-based catalyst is contacted, and the isobutene system which is in contact with the cationic part is dimerized to diisolate and hydrogenated to isooctane. This method is used for the production of isooctane. An isobutene-containing hydrocarbon feedstock FI 106955 and US 6,613,108 are contacted with an acid exchange resin in the presence of a nitrogen-containing compound, and at least one isobutene is then obtained. 1 044771 discloses a method for removing sulfur compounds from a hydrocarbon feedstock by using a hydrogenation gas in the presence of a hydrodesulfurization catalyst such as a platinum catalyst carried on an oxide catalyst to remove sulfur compounds present in the feedstock. The sulfur-containing compound is chlorinated, and the hydrocarbon feedstock and / or hydrogenation gas contains an olefin and / or carbon oxide. The method includes passing a vaporous hydrocarbon mixture and a hydrogenation gas through a hydrodesulfurization catalyst to convert existing sulfur into hydrogen sulfide, and removing the hydrogen sulfide. The product obtained by the dimerization reaction of butene typically contains a trace amount of sulfur-containing compounds as a 2004200414 substance. This ^ 4 ^. ^ ^ Dizape will affect the availability of the product isooctane as a gasoline blending component. The Santiago and Kana y > technology publications do not teach a simultaneous soaping step of hydrogenating olefins and removing sulfur content to ψ > $ Β, but we know that there is a clear need for the "Hai method." # 用 # 、 a Θ ^ In addition, the first technical publications did not teach at all the use of low hydrogen / olefins who contain a protective character ratio to hydrogenate and remove olefins at the same time [invention] An object of the present invention is to provide a ^ U is a method of blending gasoline with double < knife from paraffin hydrocarbons containing olefin feed mixture obtained from the dimerization of butene dimers; 1 and 8 + +, I 仏 3 & Human :: " The clear method includes the simultaneous hydrogenation of olefins and sulfurization by hydrogenolysis "Do not degrade." The method according to the invention requires a low excess of hydrogen. Application: The scope of the patent discloses the characteristics of the method according to the present invention, which is used for gasoline blending ingredients per sulphur-free paraffin hydrocarbon. [Embodiment] A detailed explanation of the sheep ▽ People drive different places' have found that the method can achieve the above-mentioned purpose and avoid the shortcomings of the prior art, or reduce it by the method of the present invention under specific operating conditions. According to the invention, those containing 80 to 97% by weight are used. Olefins, " 0% by weight. 12 olefins. Bu 7 weight _C9, ^ u heavier > (^ 2 thin smoke and optionally a small amount of raw materials. The raw materials can be derived from the polymerization of dilute diamine; obtained: = 200413514 compound. Raw materials 3 have 丨―1 000wt—PPm sulfur compound, typically i—50wt-ppm 'This calculation is based on sulfur, sulfide, disulfide, thiophene, and / or alkyl quinone. Suitable for use in the method of the present invention The hydrogenated catalyst is a noble metal catalyst on an alumina support. The preferred noble metal catalyst is platinum and / or palladium, particularly preferably platinum. The content of the noble metal on the support is < 1% by weight. In some cases The shell metal can be replaced by nickel. The isomerization of high octane trimethylpentene and its branched olefins in the raw material is less valuable. When alumina is used as a catalyst, the system can be suppressed. Less branched isomers. In this process, the 'feedstock is hydrogenated in two steps. In the first step, most of the C8 olefins are converted, but the conversion of heavier olefins and sulfur compounds is relatively small. In the second step, the remaining Cs olefins, olefins, and other heavier olefins and sulfur compounds Reaction. Sulfur compounds are converted into light hydrocarbons and hydrogen sulfide. In the first step, the product stream can be recycled as needed in order to dilute the concentrated olefin feed, and therefore the reaction heat is safely removed by saturation of the double bonds. The reaction temperature in the first step is in the range of 150-230 ° c, and the pressure is in the range of 20-70 bar. In the second step, a higher reaction is administered in the reactor than in the first step. Temperature. The reaction temperature is in the range of 180-300 ° C, the molding range is 1 90-260 ° C, and the pressure is in the range of 20-70 bar. In the first step, h The olefin is hydrogenated and the conversion rate is greater than 700 / 〇. 200413514 By using the relatively low reaction temperature disclosed above, the deactivation of the catalyst in the first step can be avoided or at least reduced. In the first step, A fixed-bed three-stage hydrogenation reactor is used, and the vehicle is a trickle-bed reactor. The reactor is typically used as a trickle-bed reactor to implement the trickle-bed reaction when compared with a fluid bed. Device has more than twice the fluid circulation flow rate. This control The temperature rise in the catalyst bed of the reactor. In a trickle bed reactor, the catalyst is placed in a fixed bed, while the liquid dilute feed and gaseous hydrogen feed flow down through the bed at the same time. The preferred flow patterns are trickle flow and pulsating flow patterns. Moreover, the μ motion pattern of the three-phase hydrogenation reaction is also possible, such as reverse flow and co-current flow. Use fluid circulation and by introducing fresh hydrogen into the first step to maintain The hydrogen concentration in the south of the first step. The circulating flow system is saturated with dissolved hydrogen. The advantages of the high hydrogen concentration in the first step are that it is beneficial to olefins and sulfides, and at the same time effectively reduces coke formation on the catalyst. In the process of the first step, the present invention can use one or several reactors: the reactors can be divided into more than one reactor, or the reactor can be divided into more than one reaction, with a cooler between the reactor beds. By using the interval h _ pi to reduce the% flow. The heat exchanger can be used to reverse the t-bundle material in the process, and m will react to preheat the raw materials that will enter the butene dipolymerization unit, or be used as the steaming tower of the dimer dimerization unit. The energy is directly or indirectly through the factory or the site, and the process units directly used in other sites. One step is the conversion of olefins and heavier olefins in another reactor ..., the rest of the c8 dimarite% fe, and at the same time, the remaining sulfur compounds are degraded by the hydrogenolysis 200413514 reaction. In the second step, the reaction product is not recycled, and the feed is kept as concentrated as possible. In the second step, the reactor It is also typically operated as a trickle bed reactor, but other flow methods of fixed-bed three-phase hydrogenation can be applied. In the second step, the reaction vessel can be combined with the same reaction vessel as the first reaction step if necessary. In the stabilization tower after the reactor, light hydrocarbons, hydrogen sulfide, and sulfur-containing impurities are mainly separated from the product. The present invention is illustrated in more detail in Figures 1 and 2. These diagrams only illustrate the preferred specific aspects of the present invention. Limit the scope of the invention. In Figure 1, in the first step, before entering the preliminary reactor 30, the hydrocarbon feed 10 is combined with the fresh hydrogen feed 20. The combined feeds 10 and 20 are connected After the catalyst bed 33 in the preliminary reactor 30, the gas stream 60 and the liquid stream 31 are separated from each other at the lower part of the reactor. The liquid stream 31 is cooled in a heat exchanger 80, and the heat exchanger 8G is used for cooling From the dimerization reaction of butene ^, after the procedure 8b, after cooling, ㈣32 = circulation 62 and flow 6, the circulation flow 62 returns to the combined feed ^ and other, this flow is leading to the second step, Into the second reactor, enter the second reactor 4 of the second step. Previously, the flow system was merged with the gaseous product gas stream 60 of the first step. The lower part of the reactor 4 was separated from each other, and they were connected. In the second figure, the preliminary reactor in the other step of the present invention is shown in Figure 2. ^ Λ En, where The first knife cuts into two catalyst beds with a liquid product outlet of 10 200413514 between the catalyst beds. The system is cooled and returned. The second catalyst bed. In the reactor. After the first catalyst bed, the liquid product flow can optionally be set in the reaction shown in Figure 2 in the second diagram, before the first reactor-in the second step, before the preliminary reactor 30. The smoke feed 10 is combined with the fresh hydrogen feed α. The combined feeds 10 and 20 are reacted to the first total catalyst 30 in the 30, and τ blade catalyst bed 34. In the catalyst bed 34 After that, the air flow is separated from the liquid flow, and the liquid supply Q 1, χ, ", and K is mainly used for cooling the heat exchanger 90. The cooled liquid stream 9 2 is returned to the catalyst bed 35 on the first-line, and then the gas stream 60 and the liquid stream 31 are separated from each other at the lower part of the reactor. The heat exchanger 90 can be operated like the heat exchanger 80, using a program stream from the dimerization of butene as a cooling fluid. Since the program flow 81 transmitted by Ding Xianzhi No. 2 Dendrobium No.1, the liquid flow 31 is negotiated in the heat exchanger 80; §ρ. Jealous of you Irrigation department Then, after cooling, the liquid stream 32 is divided into a circulating stream 62 and a stream β, and the earth and machine 2 are only 61. The latter leads to the second step and enters the second reactor 4: 0. Before entering Rendi Yifeng_Shi c $ into the second reactor 40 in the first step, stream 1 was combined with gas stream 60 containing gaseous products from the first step. Close = grab 6M 61 and then lead to the catalyst bed in reactor 4 (), while gas 70 and liquid flow 71 are tied together; 5 virtual cry / 1 η λα is separated from each other in the lower part of reaction benefit 40, and they pass through Go to the stabilizing towers 50, 1 and go to 4 ~ 1 ,,--###### becomes a stream 51 separate from the paraffin product stream 52. The method for producing gasoline blended components according to the present invention has several advantages when compared with today's technology. The removal of sulphur compounds and the saturation of the thin smoke occur in the same reactor under mild operating conditions τ. Use: There are less than 1% by weight of noble metal on the oxide support and hydrogenated heavy c12 olefin. The use of alumina support on precious metal catalysts will prevent isomerization of branched high-octane gasoline components into less branched low-octane components, which will be considerable when using alumina / silica support. To happen. The lower reaction temperature of the reactor in the first step reduces the deactivation of the catalyst. Furthermore, the circulation of the product stream in the reactor in the first step, as well as the circulation of intermediate beams originating from any hydrogenation reactor, ensures the safe control of temperature. In the second step, there is no circulation, and the concentrated feed ensures conversion of the remaining unreacted material. Another advantage of the present invention is that a hydrogen / olefin ratio close to the required stoichiometric consumption can be used in order to achieve the required olefin conversion and desulfurization. The hydrogen / olefin mol ratio is 0.9 to 20, preferably 105 to 15, which results in low hydrogen consumption. Low excess hydrogen and mild operating conditions will reduce the loss of hydrogen with vapor products, and it is generally not necessary to use a hydrogen cycle compressor for hydrogen recovery. The following examples further illustrate the invention, but they are not to be considered as limiting the scope of the invention. EXAMPLES Example 1 In a two-step process, the hydrogenation of olefins and the removal of sulfur compounds, and the quality of the products obtained, were used as gasoline blending ingredients in the feed. The feed was hydrogenated in a two-step process using a platinum catalyst (with 0.3% by weight palladium on an alumina support). The hydrogenation is carried out in two steps, in which a recycle stream of 6 times the fresh feed stream is applied in the first v step, and the second step does not require any recycle stream to operate. In the first step, the temperature is about 180 12 200413514 C, and in the second step is 220 ° C. In the first step, the conversion rate of octane was 92%, and the selectivity of 2,4,4-trimethylfluorene for isomerization was 98%. The feed contained 5% by weight of dodecene. 30% by weight of 2,2,6,6-tetra-4-methylene-heptane is converted, 2 '2, 4, 6, 6-pentamethyl-heptene (trans) and 2, 2, The 4, 6, 6-pentamethyl-3-heptane (cis) system was not completely transformed. In the second step, the conversion of octene was about 100% at 220 ° C, and the conversion of dodecene was 83-94%. The total olefin conversion was then 04%, and the selectivity of 4-trimethylpentene to isooctane was 100%. The sulfur content was reduced to Wm and PPm to 0.6 wt-ppm in the product. The components of the final product were analyzed using gas phase; iff, y ~ 丨, and 9 methods. The results of the analysis (PONA = paraffin, olefin, pyrene, aromatic tobacco) are shown in Table 1 below, while other results are shown in Table 2. 0
13 200413514 表3 :蒸餾結果 開始 72.1 °C 5體積％ 97.0°C 10體積％ 99.9°C 20體積％ 101.8°C 30體積％ 102.6°C 40體積％ 102.6°C 50體積％ 104.2°C 60體積％ 104.9°C 70體積％ 106.6°C 80體積％ 110.8°C 90體積％ 127.3°C 95體積％ 171.0°C 結束 205.6°C 產率 98. 2體積％ 殘餘 1.2體積％ 損失 0. 6體積％ E-70 0. 0體積％ E-100 9. 7體積％ E-150 92. 1體積％ E-180 96. 2體積％13 200413514 Table 3: Start of distillation results 72.1 ° C 5 vol% 97.0 ° C 10 vol% 99.9 ° C 20 vol% 101.8 ° C 30 vol% 102.6 ° C 40 vol% 102.6 ° C 50 vol% 104.2 ° C 60 vol% 104.9 ° C 70 vol% 106.6 ° C 80 vol% 110.8 ° C 90 vol% 127.3 ° C 95 vol% 171.0 ° C over 205.6 ° C yield 98.2 vol% residual 1.2 vol% loss 0.6 vol% E- 70 0.0% by volume E-100 9. 7% by volume E-150 92.1% by volume E-180 96.2% by volume
表2 :其它試驗結果 試驗 結果 所用的方法 硫含量 0.6 wt-ppm ASTM D-3120 R0N (辛烷值的研究） 98. 5 ASTMD-2699-97 (1998) Μ0Ν (馬達辛烷值） 96. 0 ASTM D-2700-95a (1998) 氧化安定性 >1500 min EN ISO 7536 存在和潛在的膠質 <1 mg/100 ml EN ISO 6246 RVP (雷德蒸氣壓） 20.9 kPa EN 13016 密度/15QC 708.8 kg/m3 ISO 12185 蒸餾 見表3 ISO 3405 實施例2 載體的選擇對於2, 4, 4-三甲基戊烯之影举 觸媒1 :鈀<1重量％，在ai2o3載體上 觸媒2 :鈀<1重量％，在Al 203/Si02載體上 14 200413514 進料：2, 4, 4-三甲基戊烯 條件：溫度：150°C 壓力：30巴 WHSV : 1升/小時 反應器糸統··實驗室規模的管式反應器 以下表4中顯示轉化率及選擇率。 表4 觸媒 2, 4, 4-三曱基戊烯的轉化率 觸媒1 99.9 觸媒2 98. 9 ▼ 丨 y ，w —ypu ^辛烷的 ΞΞΞΖΞΞΖΖιξ] _ 58·! 實施例3 在兩步驟中使用低過量的氩使烯烴進料氫化及硫化合 物降解 使用鉑觸媒（具有〇·3重量％的鉑在氧化鋁載體上），在 兩步驟設備中氫化由丁烯之二聚合反應所得到的^吒^烯 垣進料。在第-步驟中施^ 6倍於新鮮進料流的循環流， 而且第二步驟係不需要任何循環流來操作。在第一步驟中 ，反應器的平均操作溫度係19〇 — 198<^，且在第二步驟平 均刼作溫度為210-218°C。第一步驟中的操作壓力係2g巴 ’而第二步驟係28巴。 進料含有0.94%的C4-C16烯烴、94.0%的c锐卜 4.3〇%的C12烯烴、〇.43%的芳烴及〇 4%的其它者。進；入 有以硫計算為3wt —ppm硫化合物。氫僅導入第一牛料3 自第一階段的氫流出物係被導引至第二階段。 '^钭中氫/ 15 200413514 烯烴莫耳比係1 · 〇7 °藉由蒸流來穩定產物流，及藉由氣相 層析方法來分析它。 產物具有以下的組成.·96·9%的總石蠟、0%的C8烯烴 、2·49%的c12烯煨、0.46%的芳烴及〇·15%的其它者。心烯 烴的轉化率為10〇%。硫含量係低於檢測極限（<〇 lwppm)。 【圖式簡單說明】 第1 第2 樣 )圖式部分 圖示意地顯示依本發明的方法之較佳具體態樣。 圖不意地顯不依本發明的方法之另一較佳具體態 (二）元件代表符號 10 20 30 31 32 33 34 35 40 44 50 烴進料 新鮮的氫進料 初步反應器 液流 液流 觸媒床 第一觸媒床 第二觸媒床 第二反應器 觸媒床 穩定塔 200413514 51 流 52 石蠟產物流 60 氣流 61 流 62 循環流 70 氣流 71 液流 80 熱交換器 81 程序流 90 熱交換器 91 液流 92 液流Table 2: Other test results Test method Sulfur content 0.6 wt-ppm ASTM D-3120 R0N (octane number study) 98.5 ASTMD-2699-97 (1998) MON (motor octane number) 96.0 ASTM D-2700-95a (1998) Oxidation stability> 1500 min EN ISO 7536 Existing and potential colloids <1 mg / 100 ml EN ISO 6246 RVP (Reid vapor pressure) 20.9 kPa EN 13016 Density / 15QC 708.8 kg / m3 ISO 12185 distillation see Table 3 ISO 3405 Example 2 Selection of carrier For 2, 4, 4-trimethylpentene shadow catalyst 1: Palladium < 1% by weight, catalyst 2 on ai2o3 carrier: Palladium < 1% by weight on Al 203 / Si02 support 14 200413514 Feed: 2, 4, 4-trimethylpentene Conditions: temperature: 150 ° C pressure: 30 bar WHSV: 1 liter / hour reactor 糸System · Lab-scale tubular reactors The conversion and selectivity are shown in Table 4 below. Table 4 Conversion rate of catalyst 2, 4, 4-trimethylpentene Catalyst 1 99.9 Catalyst 2 98. 9 ▼ y, w —ypu ^ octane's ΞΞΞZOΞΞZOZO1] _ 58 ·! Example 3 In two A low excess of argon was used in the step to hydrogenate the olefin feed and degrade the sulfur compounds. A platinum catalyst (with 0.3% by weight of platinum on an alumina support) was used to hydrogenate the butylene dimerization reactor in a two-step apparatus The resulting ^ 吒 ^ ene is fed. In the first step, 6 times the circulating flow of the fresh feed stream is applied, and the second step does not require any circulating flow to operate. In the first step, the average operating temperature of the reactor is 19-198 < ^, and in the second step, the average operating temperature is 210-218 ° C. The operating pressure in the first step was 2 gbar 'and the second step was 28 bar. The feed contained 0.94% of C4-C16 olefins, 94.0% of CRIBO 4.30% of C12 olefins, 0.43% of aromatics, and 0.4% of others. Into; 3wt-ppm sulfur compounds calculated as sulfur. Hydrogen is only introduced into the first bovine feed 3. The hydrogen effluent from the first stage is directed to the second stage. '^ 钭 中 15/15 200413514 Mole ratio of olefins 1 · 07 ° The product stream is stabilized by a steam stream, and it is analyzed by gas chromatography. The product has the following composition: 96.9% of total paraffin, 0% of C8 olefins, 2.49% of C12 olefins, 0.46% of aromatics, and 0.15% of others. The conversion of cardiones was 100%. The sulfur content is below the detection limit (< 0 lwppm). [Brief description of the drawings] 1st and 2nd) Schematic part The figure schematically shows a preferred embodiment of the method according to the present invention. The figure unintentionally shows another preferred embodiment of the method according to the present invention. (II) Symbols of the elements 10 20 30 31 32 33 34 35 40 44 50 Hydrocarbon feed fresh hydrogen feed preliminary reactor liquid flow catalyst Bed first catalyst bed second catalyst bed second reactor catalyst bed stabilization tower 200413514 51 stream 52 paraffin product stream 60 gas stream 61 stream 62 circulating stream 70 gas stream 71 liquid stream 80 heat exchanger 81 process stream 90 heat exchanger 91 liquid flow 92 liquid flow
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