JP3634041B2 - Light oil quality treatment method - Google Patents
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Description
【0001】
【発明の属する技術分野】
本発明は、軽油、特に分解系軽油の高品質化処理法に関し、更に詳しくは軽油留分に相当する沸点範囲を有する留分でありながら、低いセタン価、濃い色相、難脱硫性硫黄分の高い含量等の制約から、従来、用途が限定されていた分解系軽油の高品質化処理方法に関するものである。
【0002】
【従来の技術】
ディーゼルエンジンの燃料として使用されているディーゼル燃料は、従来、原油の常圧蒸留により得られる特定の沸点範囲の直留軽油留分に水素化処理等を施して得た軽油留分をブレンドすることにより製造されている。
ところで、ディーゼル燃料の需要は近年著しく増大しているにもかかわらず、ディーゼル燃料の主基材になる軽油直留留分は、原油単位量当たり限られた量しか含まれていないために供給量が限られ、その結果、需要の増大に見合って、ディーゼル燃料を充分に供給することが難しくなる傾向にある。
【0003】
一方、接触分解装置等より得られる特定の沸点範囲の分解系軽油は、最近の白油化傾向に伴い、直留軽油留分とは逆に余剰傾向にある。そこで、分解系軽油は、軽油直留留分の不足を補う新たなブレンド基材源として注目されている。
ところで、分解系軽油には、セタン価が低く、芳香族成分の比率が高く、しかも独特の着色を呈して色相が悪いと言った品質上の問題がある。また、分解系軽油に含まれている硫黄分は、難脱硫性の化合物(例えば4,6−ジメチルジベンゾチオフェン等)が多く、深度脱硫を施して硫黄含有量を低減することが困難である。
これらの問題点を解決して分解系軽油を高品質化し、商品軽油として、又は軽油のブレンド基材として利用する試みが成されているが、以下に説明するような理由から実用化されるには到っていない。
【0004】
例えば、試みの一つとして、分解系軽油を水素化して有効利用する方法が、特開昭63−291985号公報に開示されている。しかし、分解系軽油を単独で商品軽油とするためには、色相及びセタン価を向上させることが必要であるが、前掲公報による方法では、コストが嵩み、経済的に引き合わない。
すなわち、この方法によりセタン価を向上させるためには水素化反応の温度を高くする必要があるが、温度を高くすると、色相が著しく劣化して、ブレンド基材として不適当になる。また、水素化反応の温度を高くする代わりに水素化反応の圧力(水素分圧)を高くすると、反応装置の耐圧強度を上げる必要が生じて設備費及び運転費が嵩み、ブレンド基材の価格が上昇し、コスト的に非常に不利になる。
【0005】
また、別法として、分解系軽油を軽油直留留分にブレンドし、脱硫処理と水素化処理からなる二段処理を行って高品質化し、分解系軽油を有効利用する方法が、特願平6−339184号公報に開示されている。
しかし、分解系軽油を直留軽油留分にブレンドした原料油中には難脱硫性硫黄化合物が多く含まれるために、深度脱硫する際、反応温度を高くする必要がある。反応温度を高くすると、生成油が著しく着色する。また、白金、ロジウム、ルテニウム系の水素化能が高い触媒を使用して水素化処理を行うと、触媒が硫黄被毒され活性を喪失し易いために、実用上必要な触媒寿命を確保できないと言う問題がある。
【0006】
【発明が解決しようとする課題】
以上のように、従来の方法では、分解系軽油を高品質化処理し、ディーゼル燃料のブレンド基材として、また商品軽油として活用することが難しい。
そこで、本発明は、従来技術の有する上記問題点を克服して、原料軽油を高品質化処理し、セタン価が高く、色相に優れ、しかも硫黄含量の低い高品質な軽油留分に転化する方法を提供することを目的としている。
【0007】
【課題を解決するための手段】
本発明者らは、上記の目的を達成するために検討を重ねた結果、接触反応装置等で製造される分解系留出油を特定の蒸留条件下で蒸留し、得た特定沸点範囲の分解系軽油留分を特定の反応条件下で脱硫し、得た水素化脱硫生成物を特定の反応条件下で更に水素化処理を行うことによって得られる軽油留分が、セタン価及び色相に優れ、硫黄含有量の少ないと言う優れた性状を有することを見い出し、本発明を完成するに至った。
【0008】
上記目的を達成するために、本発明に係る軽油の高品質化処理方法は、3〜8MPaの範囲の水素分圧、220〜380℃の範囲の反応温度、1.0〜5.0hr−1の範囲の液空間速度及び100〜400L/Lの範囲の水素/オイル比の反応条件下で無機酸化物からなる担体に周期律表6族金属および8族金属を担持させた脱硫触媒と沸点範囲160〜340℃であって、かつ沸点範囲の上限が320℃以上340℃以下の軽油留分とを接触反応させて軽油留分を水素化脱硫する脱硫工程と、
白金、ロジウム、ルテニウムから選ばれた少なくとも1種類以上の金属を含んでなる水素化触媒、又はニッケル及びタングステンを含んでなる水素化触媒の存在下で3〜8MPaの範囲の水素分圧、200〜350℃の範囲の反応温度、1.0〜5.0hr−1の範囲の液空間速度、100〜400L/Lの範囲の水素/オイル比の反応条件で前記水素化脱硫工程を経た軽油留分を水素化処理して、セタン価が45〜53、セーボルト色値が2〜12、かつ1環芳香族含有量が13.0〜29.0容積%、飽和分含有量が69.5〜86.5容積%、2環芳香族含有量が0.5〜2.2容積%の生成油を生成するようにした水素化処理工程と
を備えていることを特徴としている。
【0009】
本発明方法で原料油とする軽油留分は、沸点範囲が160〜340℃、好ましくは160〜330℃、さらに好ましくは160〜320℃の軽油留分である。本発明方法では、例えば接触分解軽油、熱分解軽油、直留軽油、水素化処理軽油及び脱硫処理軽油を単独で原料油としても良く、それらを混合して原料油としても良い。
【0010】
上記軽油留分の水素化脱硫条件において、軽油留分の沸点範囲の上限が340℃より高いと、難脱硫性の硫黄化合物、例えば4,6−ジメチルジベンゾチオフェン等の量が著しく増加し、こうした原料軽油留分を深度脱硫するために、反応温度を高くする必要が生じ、生成油の色相が悪化する。逆に、沸点範囲の上限が320℃未満であると、生成油のセタン価が著しく低下するので、原料油として好ましくない。
【0011】
また、水素化脱硫触媒の負荷を軽減する意味から、水素化脱硫工程に導入する軽油留分は、その硫黄含有量が少ない方が好ましく、1質量%以下が好適である。予め、原料軽油留分を蒸留操作によって特定範囲の沸点を持つ留分、すなわち、沸点範囲160〜340℃、好ましくは160〜330℃、さらに好ましくは160〜320℃の留分として分留することにより、軽油留分中の硫黄分を1質量%以下にすることができる。
【0012】
水素化脱硫工程は、軽油留分の硫黄分の除去を主な目的として脱硫触媒により水素化脱硫処理する。脱硫工程の反応条件は、水素分圧が3〜8MPa、好ましくは4〜7MPaであり、温度が220〜380℃、好ましくは250〜350℃であり、液空間速度が1.0〜5.0hr−1、好ましくは1.0〜3.0hr−1であり、水素/オイル比が100〜400L/L、好ましくは200〜300L/Lである。
【0013】
圧力(水素分圧)が3MPa未満であると、触媒の脱硫活性が低下すると共に生成油の色相も悪化し、逆に8MPaを超えると、設備の耐圧を高くすために設備費と運転費が嵩み、不経済になる。
反応温度が220℃未満であると、触媒の脱硫活性が低く、380℃を超えると、脱硫活性が飽和するために温度を上げても脱硫効果が向上しないばかりでなく、生成油の色相が悪化し、しかも設備費と運転費が嵩む。
液空間速度が5.0hr−1を超えると、触媒と原料油の接触時間が短くなり過ぎ、脱硫反応が十分に行われないために生成油の残留硫黄分が多くなり、1.0hr−1未満では必要以上に接触時間が長くなり過ぎ、処理効率が低下する。
水素/オイル比が100L/L未満であると、十分に脱硫反応が進まず、逆に400L/Lを超えると、過剰の水素を消費することになるので、処理コストが増大し不経済である。
【0014】
本発明方法で使用する水素化脱硫触媒は、担体として、種々のものが使用できる。例えば、担体として、シリカ、アルミナ、ボリア、マグネシア、チタニア、シリカ−アルミナ、シリカ−マグネシア、シリカ−ジルコニア、シリカ−トリア、シリカ−ベリリア、シリカ−チタニア、シリカ−ボリア、アルミナ−ジルコニア、アルミナ−チタニア、アルミナ−ボリア、アルミナ−クロミア、チタニア−ジルコニア、シリカ−アルミナ−トリア、シリカ−アルミナ−ジルコニア、シリカ−アルミナ−マグネシア、シリカ−マグネシア−ジルコニアなどの無機酸化物、またはこれらの1種以上の混合物が挙げられる。
これらの無機酸化物のうち、好ましいものとしては、アルミナ、シリカ−アルミナ、アルミナ−ボリア、アルミナ−ジルコニアが挙げられ、特に好ましいものとしては、アルミナ、シリカ−アルミナ、アルミナ−チタニア、アルミア−ボリア、アルミナ−ジルコニアが挙げられ、特に好ましくはγ−アルミナが挙げられる。
これらの無機酸化物は、1種類を単独で用いてもよいし、2種類以上を組み合わせて用いてもよい。
【0015】
本発明方法で使用する水素化脱硫触媒では、担体に担持させる6族金属は、触媒基準かつ酸化物換算で10〜25質量%の範囲の含有量で含まれるモリブデン、又は触媒基準かつ酸化物換算で0.1〜5質量%の範囲の含有量で含まれるタングステンのいずれかである。
モリブデンの含有量が10質量%より少ないと、活性点として働くモリブデンの絶対量が少ないために、脱硫活性が発現せず、逆に担持させるモリブデンの含有量が25質量%より多過ぎると、金属の凝集が起こり活性点の数が減少し、その結果、脱硫活性が却って低下する。
タングステンの含有量が0.1質量%より少ないと、難脱硫性物質の核水素化活性が発現せず、逆に5質量%より多いと、担持するタングステンの量が多すぎ、金属の凝集が起こり、その結果、脱硫活性が却って低下する。
【0016】
また、本発明方法で使用する水素化脱硫触媒で担体に担持させる8族金属は、触媒基準かつ酸化物換算で1〜10質量%の範囲の含有量の範囲で含まれるコバルト及びニッケルの少なくとも一方である。
8族金属の含有量が、1質量%より少ないと充分な脱硫活性が得られず、また10質量%を超えると、脱硫活性が飽和し、それ以上の添加効果が見い出せなくなる。
更に、必要に応じて、6族金属及び8族金属からなる活性金属に加えて、リン、ホウ素、亜鉛、ジルコニア等を含ませることができる。
【0017】
完成触媒の平均細孔径は、60〜90Åであることが好ましい。平均細孔径が大きすぎると、細孔内への硫黄化合物の拡散性は良いものの、触媒の有効表面積が小さくなるので、脱硫活性が低下する。
【0018】
本発明方法では、水素化脱硫工程を経た軽油留分は、更に水素化処理工程で水素化処理される。水素化処理工程の反応条件は、圧力(水素分圧)が3〜8MPaの範囲、好ましくは4〜7MPaの範囲、反応温度が200〜360℃の範囲、好ましくは250〜350℃の範囲、液空間速度が1.0〜5.0hr−1の範囲、好ましくは1.0〜3.0hr−1の範囲、水素/オイル比が100〜400L/Lの範囲、好ましくは200〜300L/Lの範囲にある。
【0019】
圧力(水素分圧)が3MPa未満であると、水素化反応が十分に行われないのみならず、セタン価の向上が見られず、色相も悪化し、逆に8MPaを超えると、設備の耐圧を高くすために設備費と運転費が嵩み、不経済になる。
反応温度が200℃未満であると、触媒の水素化活性が低く、360℃を超えると、水素化活性が飽和するために温度を上げても水素化処理効果が向上しないばかりでなく、生成油の色相が却って悪化し、しかも設備費と運転費が嵩む。
液空間速度が5.0hr−1を超えると、触媒と原料油の接触時間が短くなりすぎ、水素化反応が十分に行われず、1.0hr−1未満では必要以上に接触時間が長くなりすぎ、処理効率が低下する。
水素/オイル比が100L/L未満であると、十分な水素化反応が得られず、逆に400を超えると、過剰の水素を消費することになるので処理コストが増大し不経済である。
【0020】
本発明方法で使用する水素化処理触媒は、担体として種々のものが使用できる例えば、担体として、シリカ、アルミナ、ボリア、マグネシア、チタニア、シリカ−アルミナ、シリカ−マグネシア、シリカ−ジルコニア、シリカ−トリア、シリカ−ベリリア、シリカ−チタニア、シリカ−ボリア、アルミナ−ジルコニア、アルミナ−チタニア、アルミナ−ボリア、アルミナ−クロミア、チタニア−ジルコニア、シリカ−アルミナ−トリア、シリカ−アルミナ−ジルコニア、シリカ−アルミナ−マグネシア、シリカ−マグネシア−ジルコニアなどの無機酸化物、またはこれらの2種以上の混合物が挙げられる。
これらの無機酸化物のうち、好ましいものとしては、アルミナ、シリカ−アルミナ、アルミナ−ボリア、アルミナ−ジルコニアが挙げられ、特に好ましいものとしては、アルミナ、シリカ−アルミナ、アルミナ−チタニア、アルミア−ボリア、アルミナ−ジルコニアが挙げられ、特に好ましくはγ−アルミナが挙げられる。
これらの無機酸化物は、1種類単独で用いてもよいし、2種類以上を組み合わせて用いてもよい。
【0021】
本発明方法で使用する水素化処理触媒に含有させる水素化活性金属は、白金、ロジウム、ルテニウム等が使用でき、それらの1種以上を併用しても差し支えない。
これらの水素化活性金属の含有量は0.1から3質量%とするのが良い。0.1質量%未満とすると、水素化活性が充分でなく、逆に3質量%を超えると、担持量を増やしても水素化活性の向上が期待できない。
また、上記の活性金属に代えてニッケル及びタングステンを使用することもできる。この場合のニッケルの含有量は1〜10質量%とし、タングステン含有量は0.1〜5質量%とするのが良い。
必要に応じて、これらの活性金属に加えて、リン、ホウ素、亜鉛、ジルコニア等を含ませることができる。
【0022】
上述のように、本発明に係る分解系軽油の高品質化処理法は、特定の沸点範囲の軽油、例えば分解系軽油留分を特定の条件で水素化脱硫した後、特定の条件でさらに水素化処理することにより、高いセタン価と良好なカラーを有し、しかも残留硫黄分の低い軽油留分として高品質化する方法である。尚、水素化脱硫工程と水素化処理工程との間には、既知の硫化水素除去装置を設けて水素化脱硫工程で生じた硫化水素を除去することが好ましい。
また、本発明方法は、高圧流通式反応器に固定床式触媒を形成した通常の反応装置を使用し、かつ比較的マイルドな反応条件で軽油留分を処理できる。よって、本発明方法によれば、例えば余剰の分解系軽油をディーゼル燃料等の軽油留分に経済的に転化することができる。
【0023】
【発明の実施の形態】
以下に、実施例を挙げ、添付図面を参照して、本発明方法の実施の形態を具体的かつ詳細に説明する。
図1は、本発明方法を実施する反応装置の一例の構成を示す模式図である。図1に示す反応装置10は、高圧流通式反応装置であって、水素化脱硫工程を実施する反応器12と、水素化処理工程を実施する反応器14とが直列に接続されている。反応器12には、無機酸化物からなる担体に周期律表6族金属および8族金属を担持させた水素化脱硫触媒の固定床式触媒層16が形成され、一方、反応器14には、白金、ロジウム、ルテニウムから選ばれる少なくとも1種類以上の金属を含んでなる水素化触媒、又はニッケル及びタングステンを含んでなる水素化処理触媒の固定床式触媒層18が形成されている。
反応器12と14との間に、反応器12から流出した水素化脱硫生成物の内、硫化水素を取り除くための既知の硫化水素除去装置20が設けられている。また、反応器14の下流には、反応器14の下部から流出した生成物を生成油とガスに分離するための気液分離槽22が設けてある。
【0024】
沸点範囲160〜340℃の軽油留分からなる原料油と水素含有ガスとの混合流体は、加熱炉及び熱交換器等(図示せず)により反応温度まで加熱されて、反応器12の上部から反応器12に導入される。軽油留分は、反応器12内で、3〜8MPaの範囲の水素分圧、220〜380℃の範囲の反応温度、1.0〜5.0hr−1の範囲の液空間速度及び100〜400L/Lの範囲の水素/オイル比の反応条件下で固定床式触媒層16の水素化脱硫触媒に接触して水素化脱硫される。
【0025】
水素化脱硫工程を経た軽油留分は、必要に応じて更に水素含有ガスが注入され、また反応温度にまで加熱又は冷却されて反応器14の上部から反応器14に導入される。軽油留分は、反応器14内で、3〜8MPaの範囲の水素分圧、200〜360℃の範囲の反応温度、1.0〜5.0hr−1の範囲の液空間速度、100〜400L/Lの範囲の水素/オイル比の反応条件で固定床式触媒層18の水素化処理触媒に接触して水素化処理される。反応器14の下部から流出した生成物は、気液分離槽20で生成油とガスとに気液分離される。通常、反応器12と反応器14との間に圧縮機を設ける必要がないように、水素化処理工程は、水素化脱硫工程の運転圧力より多少低い運転圧力で運転される。尚、図1は本発明方法を適用する高圧流通式反応装置の構成を示す模式図であって、加熱炉、熱交換器等の本発明方法と直接関連しない機器は省略されている。
【0026】
【実施例】
実施例1
本発明方法を評価するために、接触分解装置から得た、以下に示す性状の沸点範囲160〜320℃の分解系軽油を原料にして、図1に示す高圧流通式反応装置10と同じ構成の実験装置によって本発明に係る水素化脱硫処理工程及び水素化処理工程を実施した。
【0027】
原料油(沸点範囲160〜320℃の分解系軽油)の性状
比重(15/4℃) :0.9046
硫黄分 :0.05質量%
窒素分 :0.021質量%
蒸留性状
沸点(IBP) :160℃
沸点(50%点) :254.5℃
沸点(FBP) :320℃
粘度(@30℃) :2.641mm2 /s
組成
飽和炭化水素 :28容積%
1環芳香族化合物 :32容積%
2環芳香族化合物 :38容積%
3環以上芳香族化合物 :2容積%
難脱硫性化合物
(4,6−ジメチルジベンゾチオフェン):0.002質量%
【0028】
先ず、水素化脱硫触媒としてCo−Mo−W/Al2 O3 触媒(CoO:MoO3 :WO3 =5:25:3.8(質量%))を、及び水素化処理触媒としてPt/Al2 O3 触媒(Pt=1質量%)をそれぞれ図1の高圧流通式反応装置10の反応器12及び反応器14に充填して、固定床式触媒層16及び固定床式触媒層18を形成し、それぞれ触媒に下記の前処理条件で前処理を施した。尚、この時に使用したCo−Mo−W/Al2 O3 触媒の平均細孔径は82Åであった。
【0029】
水素化脱硫触媒の前処理条件
圧力(水素分圧):3.5MPa
雰囲気:硫化水素/水素の混合ガス流通下
温度:ステップ昇温
100℃にて2hr
250℃にて2hr
350℃にて2hr
【0030】
水素化処理触媒の前処理条件
圧力(水素分圧):3.5MPa
雰囲気:水素のガス流通下
温度:ステップ昇温
100℃にて2hr
250℃にて2hr
350℃にて2hr
【0031】
次いで、反応温度に加熱した原料油と水素含有ガスとの混合流体を反応器12の上部より反応器12に導入して、以下の条件で水素化脱硫工程を実施した。更に、水素化脱硫工程を経た軽油留分とガスの混合流体を更に反応器14の上部から反応器14に導入し、以下の条件で水素化処理工程を実施し、生成油とガスの混合流体を反応器12の下部より流出させ、気液分離器20で生成油を分離した。
【0032】
水素化脱硫工程の反応条件
温度 :300℃
圧力(水素分圧):5MPa
液空間速度 :1.5hr−1
水素/オイル比 :200L/L
【0033】
水素化処理工程の反応条件
温度 :300℃
圧力(水素分圧):5MPa
液空間速度 :1.5hr−1
水素/オイル比 :200L/L
【0034】
水素化脱硫工程を経た生成油の組成、及び水素化処理工程を経た生成油の組成及び性状は、それぞれ表1及び表2に示す通りである。表2では、カラーは、セイボルト色で表示されている。
表2において、硫黄分が低く、セタン価の数値が高く、カラーの数値が大きいほど、ディーゼル燃料のブレンド基材として好ましいと評価できる。
【表1】
【表2】
【0035】
実施例2
接触分解装置から得た、以下に示す性状の沸点範囲160〜330℃の分解系軽油を原料としたこと以外は、実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
水素化脱硫工程を経た生成油及び水素化処理工程を経た生成油の組成及び性状は、それぞれ表1及び表2に示す通りである。尚、以下の実施例3から実施例12及び比較例1から比較例4の水素化脱硫工程を経た生成油及び水素化処理工程を経た生成油の組成及び性状は、実施例2と同様に、それぞれ表1及び表2に示される。
【0036】
原料油(沸点範囲160〜330℃の分解系軽油)の性状
比重(15/4℃) :0.9095
硫黄分 :0.09質量%
窒素分 :0.041質量%
蒸留性状
沸点(IBP) :160℃
沸点(50%点) :263℃
沸点(FBP) :330℃
粘度(@30℃) :2.821mm2 /s
組成
飽和炭化水素 :28容積%
1環芳香族化合物 :30容積%
2環芳香族化合物 :38容積%
3環以上芳香族化合物 :4容積%
難脱硫性化合物
(4,6−ジメチルジベンゾチオフェン):0.01質量%
【0037】
実施例3
接触分解装置から得た、以下に示す性状の沸点範囲160〜340℃の分解系軽油を原料としたこと以外は、実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
【0038】
原料油(沸点範囲160〜340℃の分解系軽油)の性状
比重(15/4℃) :0.9143
硫黄分 :0.13質量%
窒素分 :0.053質量%
蒸留性状
沸点(IBP) :160℃
沸点(50%点) :270℃
沸点(FBP) :340℃
粘度(@30℃) :3.105mm2 /s
組成
飽和炭化水素 :28容積%
1環芳香族化合物 :26容積%
2環芳香族化合物 :39容積%
3環以上芳香族化合物 :7容積%
難脱硫性化合物
(4,6−ジメチルジベンゾチオフェン):0.015質量%
【0039】
実施例4
水素化脱硫工程の反応温度を実施例1の300℃に代えて250℃としたこと以外は、実施例1と同じ分解系軽油に実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
実施例5
水素化脱硫工程の圧力(水素分圧)及び水素/オイル比を実施例1の5MPa及び200L/Lに代えてそれぞれ6MPa及び100L/Lとしたこと以外は、実施例1と同じ分解系軽油に実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
実施例6
水素化脱硫工程の反応温度及び液空間速度を実施例1の300℃及び1.5hr−1に代えてそれぞれ350℃及び4.0hr−1としたこと以外は、実施例1と同じ分解系軽油に実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
【0040】
実施例7
水素化脱硫工程の圧力(水素分圧)を実施例2の5MPa(即ち、実施例1と同じ圧力)に代えて6MPaとしたこと以外は、実施例2と同じ分解系軽油に実施例2と同じ条件(即ち、実施例1と同じ条件)で水素化脱硫工程及び水素化処理工程を実施した。
【0041】
実施例8
水素化脱硫工程の圧力(水素分圧)を実施例3の5MPa(即ち、実施例1と同じ圧力)に代えて7MPaとしたこと以外は、実施例3と同じ分解系軽油に実施例3と同じ条件(即ち、実施例1と同じ条件)で水素化脱硫工程及び水素化処理工程を実施した。
【0042】
実施例9
水素化脱硫工程で使用した触媒を実施例1のCo−Mo−W/Al2 O3 (CoO:MoO3 :WO3 =5:25:3.8(質量%))に代えてCo−Mo/Al2 O3 (CoO:MoO3 =5:20(質量%))としたこと以外は、実施例1と同じ分解系軽油に実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
【0043】
実施例10
水素化脱硫工程で使用した触媒を実施例2のCo−Mo−W/Al2 O3 (CoO:MoO3 :WO3 =5:25:3.8(質量%))(即ち、実施例1と同じ触媒)に代えてCo−Mo/Al2 O3 (CoO:MoO3 =5:20(質量%))としたこと以外は、実施例2と同じ分解系軽油に実施例2と同じ条件(即ち、実施例1と同じ条件)で水素化脱硫工程及び水素化処理工程を実施した。
【0044】
実施例11
水素化脱硫工程で使用した触媒を実施例3のCo−Mo−W/Al2 O3 (CoO:MoO3 :WO3 =5:25:3.8(質量%))(即ち、実施例1と同じ触媒)に代えてCo−Mo/Al2 O3 (CoO:MoO3 =5:20(質量%))としたこと以外は、実施例3と同じ分解系軽油に実施例3と同じ条件(即ち、実施例1と同じ条件)で水素化脱硫工程及び水素化処理工程を実施した。
【0045】
比較例5
水素化処理工程に使用した触媒を実施例1のPt/Al2 O3 (Pt=1質量%)に代えてNi−W/Al2 O3 (NiO:WO3 =5:20(質量%))とし、水素化処理工程の反応温度を実施例1の300℃に代えて360℃としたこと以外は、実施例1と同じ分解系軽油に実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
【0046】
実施例13
水素化処理工程の圧力(水素分圧)を実施例1の5MPaに代えて3MPaとしたこと以外は、実施例1と同じ分解系軽油に実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
【0047】
比較例1
接触分解装置から得た、以下に示す性状の沸点範囲160〜360℃の分解系軽油を原料とし、水素化脱硫処理に使用した触媒を実施例1のCo−Mo−W/Al2 O3 (CoO:MoO3 :WO3 =5:25:3.8(質量%))に代えてCo−Mo/Al2 O3 (CoO:MoO3 =3:15(質量%))としたこと以外は実施例1と同様の方法で行った。
【0048】
原料油(沸点範囲160〜360℃の分解系軽油)の性状
比重(15/4℃) :0.9241
硫黄分 :0.18質量%
窒素分 :0.059質量%
蒸留性状
沸点(IBP) :160℃
沸点(50%点) :281.5℃
沸点(FBP) :360℃
粘度(@30℃) :3.908mm2 /s
組成
飽和炭化水素 :28容積%
1環芳香族化合物 :24容積%
2環芳香族化合物 :37容積%
3環以上芳香族化合物 :11容積%
難脱硫性化合物
(4,6−ジメチルジベンゾチオフェン):0.025質量%
【0049】
比較例2
以下に示す性状を有する沸点範囲140〜300℃の分解系軽油を原料とし、水素化脱硫工程で使用した触媒を実施例1のCo−Mo−W/Al2 O3 (CoO:MoO3 :WO3 =5:25:3.8(質量%))に代えてCo−Mo/Al2 O3 (CoO:MoO3 =3:15(質量%))にしたこと以外は、実施例1と同じ条件で、水素化脱硫工程及び水素化処理工程を実施した。
【0050】
原料油(沸点範囲140〜300℃の分解系軽油)の性状
比重(15/4℃) :0.9003
硫黄分 :0.04質量%
窒素分 :0.02質量%
蒸留性状
沸点(IBP) :140℃
沸点(50%点) :236℃
沸点(FBP) :300℃
粘度(@30℃) :2.498mm2 /s
組成
飽和炭化水素 :30容積%
1環芳香族化合物 :34容積%
2環芳香族化合物 :35容積%
3環以上芳香族化合物 :1容積%
難脱硫性化合物
(4,6−ジメチルジベンゾチオフェン):0.001質量%
【0051】
比較例3
実施例3と同じ沸点範囲160〜340℃の分解系軽油を原料とし、水素化脱硫処理に使用した触媒を実施例1のCo−Mo−W/Al2 O3 (CoO:MoO3 :WO3 =5:25:3.8(質量%))に代えてCo−Mo/Al2 O3 (CoO:MoO3 =3:15(質量%))とし、水素化脱硫処理の反応温度を実施例1の300℃に代えて390℃としたこと以外は、実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
【0052】
比較例4
接触分解装置から得た沸点範囲160〜360℃の分解系軽油を原料としたこと以外は、実施例1と同じ条件で水素化脱硫工程及び水素化処理工程を実施した。
【0053】
表2に示すように、本発明方法で規定した条件の下で所定の分解系軽油を処理した実施例1から11及び実施例13は、残留硫黄分が0.4質量%以下で、セタン価が高く、カラーが優れている軽油留分に分解系軽油を転化することができる。
【0054】
一方、表2において、沸点範囲が160〜360℃で本発明で規定している沸点範囲160〜340℃の上限より上限が高い比較例1は、本発明方法で規定する沸点範囲内の原料油を処理した実施例1〜11及び13に比べて、セタン価及びカラーが遙に劣り、また比較例1の芳香族化合物の含量は実施例1〜11及び13に比べて遙に多い。
また、沸点範囲が140〜360℃で本発明で規定している沸点範囲の下限より下限が低い比較例2は、実施例1〜11及び13に比べて、セタン価及びカラーが遙に劣る。
また、水素化脱硫工程の反応温度が390℃で本発明で規定する反応温度の範囲を超えている比較例3は、本発明方法で規定する水素化脱硫工程の反応温度の範囲220〜380で処理した実施例1〜11及び13に比べて、特にカラーが著しく劣っている。
更に、沸点範囲が160〜360℃で本発明で規定している沸点範囲の上限より上限が高い比較例4は、実施例1〜11及び13に比べて、カラーが著しく劣り、また比較例1の芳香族化合物の含量は実施例1〜11及び13に比べて遙に多い。
比較例5は、水素化処理工程の反応温度を実施例1の300℃に代えて360℃に変えている。その結果、本発明で規定している反応温度の範囲200〜350℃を超えるために、実施例1〜11及び13に比べてセタン価が低く、1環芳香族含有量が高く、飽和分容積%が低い。
【0055】
【発明の効果】
本発明の分解系軽油の高品質化処理方法によれば、特別な分離装置や反応器を必要とせず、通常の反応装置を用いて、比較的マイルドな処理条件で特定の沸点範囲の軽油を特定の条件で処理することにより、セタン価が高く、色相に優れ、しかも硫黄含有量の少ない高品質な軽油留分を得ることができる。
これにより、余剰の分解系軽油を活用して、不足するディーゼル燃料を経済的に製造することができる。
【図面の簡単な説明】
【図1】高圧流通式反応装置の構成を示す模式図である。
【符号の説明】
10 高圧流通式反応装置
12、14 反応器
16、18 固定床式触媒層
20 硫化水素除去装置
22 気液分離器[0001]
BACKGROUND OF THE INVENTION
The present invention relates to a method for improving the quality of light oil, particularly cracked light oil, and more specifically, a fraction having a boiling range corresponding to a light oil fraction, but having a low cetane number, a dark hue, a hardly desulfurizing sulfur content. The present invention relates to a method for improving the quality of cracked gas oil, which has heretofore been limited in use because of its high content.
[0002]
[Prior art]
Diesel fuels used as fuel for diesel engines are blended with light oil fractions obtained by subjecting straight-run gas oil fractions of a specific boiling range obtained by atmospheric distillation of crude oil to hydrotreating, etc. It is manufactured by.
By the way, although the demand for diesel fuel has increased remarkably in recent years, the amount of gas oil straight cut, which is the main base material for diesel fuel, contains only a limited amount per unit of crude oil. As a result, there is a tendency that it is difficult to sufficiently supply diesel fuel to meet the increase in demand.
[0003]
On the other hand, cracked diesel oil in a specific boiling range obtained from a catalytic cracker or the like tends to have a surplus, contrary to the straight-run diesel fuel fraction, with the recent trend toward white oil. Accordingly, cracked diesel oil has attracted attention as a new blend base material source that compensates for the shortage of gas oil straight fraction.
By the way, cracked light oil has a quality problem that it has a low cetane number, a high ratio of aromatic components, and a unique coloration and a poor hue. In addition, the sulfur content contained in the cracked light oil is often difficult to desulfurize compounds (for example, 4,6-dimethyldibenzothiophene), and it is difficult to reduce the sulfur content by performing deep desulfurization.
Attempts have been made to solve these problems, improve the quality of cracked diesel oil, and use it as a commercial diesel oil or as a blend base for diesel oil, but it will be put to practical use for the reasons explained below. Is not reached.
[0004]
For example, as one of the trials, a method for effectively utilizing hydrogenated cracked gas oil is disclosed in JP-A-63-291985. However, it is necessary to improve the hue and cetane number in order to make the cracked light oil alone into commercial light oil, but the method according to the above publication is costly and is not economically competitive.
That is, in order to improve the cetane number by this method, it is necessary to increase the temperature of the hydrogenation reaction. However, when the temperature is increased, the hue is remarkably deteriorated and becomes unsuitable as a blend base material. In addition, if the hydrogenation reaction pressure (hydrogen partial pressure) is increased instead of increasing the temperature of the hydrogenation reaction, it is necessary to increase the pressure resistance of the reactor, increasing the equipment cost and operating cost, The price rises and the cost becomes very disadvantageous.
[0005]
Another method is to blend cracked gas oil into a straight fraction of light oil and improve the quality by performing a two-stage process consisting of desulfurization and hydrogenation to effectively use cracked gas oil. No. 6-339184.
However, since the raw material oil obtained by blending the cracked diesel oil with the straight-run diesel oil fraction contains a large amount of hardly desulfurizable sulfur compounds, it is necessary to increase the reaction temperature during the deep desulfurization. When the reaction temperature is increased, the product oil is markedly colored. In addition, when a hydrogenation treatment is performed using a platinum, rhodium, or ruthenium-based catalyst having a high hydrogenation ability, the catalyst is easily poisoned with sulfur and loses its activity. There is a problem to say.
[0006]
[Problems to be solved by the invention]
As described above, in the conventional method, it is difficult to improve the quality of the cracked diesel oil and use it as a blend base material for diesel fuel or as a commercial diesel oil.
Therefore, the present invention overcomes the above-mentioned problems of the prior art, converts the raw gas oil into a high-quality gas oil fraction having a high cetane number, excellent hue, and low sulfur content. It aims to provide a method.
[0007]
[Means for Solving the Problems]
As a result of repeated studies to achieve the above-mentioned object, the present inventors distilled a cracked distillate produced in a catalytic reactor or the like under specific distillation conditions, and obtained a decomposition in a specific boiling range. A gas oil fraction obtained by desulfurizing a diesel oil fraction under specific reaction conditions and further hydrotreating the obtained hydrodesulfurization product under specific reaction conditions is excellent in cetane number and hue, It has been found that it has an excellent property of low sulfur content, and the present invention has been completed.
[0008]
In order to achieve the above-mentioned object, the method for improving the quality of light oil according to the present invention comprises a hydrogen partial pressure in the range of 3 to 8 MPa, a reaction temperature in the range of 220 to 380 ° C., and 1.0 to 5.0 hr.-1Catalyst having a liquid space velocity in the range of 100 and a hydrogen / oil ratio in the range of 100 to 400 L / L and a desulfurization catalyst in which a group 6 metal and a group 8 metal are supported on a carrier made of an inorganic oxide and a boiling point range A desulfurization step of hydrodesulfurizing the gas oil fraction by contact reaction with a gas oil fraction having a boiling point of 160 to 340 ° C and an upper limit of the boiling point range of 320 ° C to 340 ° C;
Hydrogen partial pressure in the range of 3-8 MPa in the presence of a hydrogenation catalyst comprising at least one metal selected from platinum, rhodium, ruthenium, or a hydrogenation catalyst comprising nickel and tungsten, 200- Reaction temperature in the range of 350 ° C., 1.0-5.0 hr-1The hydrodesulfurization step is carried out under reaction conditions of a liquid space velocity in the range of 100 to 400 L / L and a hydrogen / oil ratio in the range of 100 to 400 L / L.PassedThe gas oil fraction is hydrotreated and the cetane number is45-53, Savort color value is2-12And the monocyclic aromatic content is13.0-29.0Volume%, saturation content is69.5-86.5Volume%, bicyclic aromatic content is0.5-2.2A hydrotreating process adapted to produce a volume% product oil;
It is characterized by having.
[0009]
The light oil fraction used as the feedstock in the method of the present invention is a light oil fraction having a boiling point range of 160 to 340 ° C, preferably 160 to 330 ° C, more preferably 160 to 320 ° C. In the method of the present invention, for example, catalytic cracking gas oil, pyrolysis gas oil, straight-run gas oil, hydrotreated gas oil, and desulfurized gas oil may be used alone, or they may be mixed to form a feed oil.
[0010]
In the hydrodesulfurization conditions of the gas oil fraction, if the upper limit of the boiling range of the gas oil fraction is higher than 340 ° C., the amount of a difficult-to-desulfurize sulfur compound such as 4,6-dimethyldibenzothiophene is remarkably increased. In order to deeply desulfurize the gas diesel oil fraction, it is necessary to increase the reaction temperature, and the hue of the product oil deteriorates. On the other hand, if the upper limit of the boiling range is less than 320 ° C., the cetane number of the product oil is remarkably lowered, which is not preferable as a raw material oil.
[0011]
Further, in order to reduce the load of the hydrodesulfurization catalyst, the gas oil fraction introduced into the hydrodesulfurization step preferably has a lower sulfur content, and preferably 1% by mass or less. The raw gas oil fraction is preliminarily fractionated as a fraction having a boiling point in a specific range by distillation operation, that is, a fraction having a boiling range of 160 to 340 ° C, preferably 160 to 330 ° C, more preferably 160 to 320 ° C. Thus, the sulfur content in the gas oil fraction can be reduced to 1% by mass or less.
[0012]
In the hydrodesulfurization step, hydrodesulfurization treatment is performed with a desulfurization catalyst mainly for the purpose of removing the sulfur content of the gas oil fraction. The reaction conditions of the desulfurization step are a hydrogen partial pressure of 3 to 8 MPa, preferably 4 to 7 MPa, a temperature of 220 to 380 ° C., preferably 250 to 350 ° C., and a liquid space velocity of 1.0 to 5.0 hr.-1, Preferably 1.0 to 3.0 hr-1The hydrogen / oil ratio is 100 to 400 L / L, preferably 200 to 300 L / L.
[0013]
When the pressure (hydrogen partial pressure) is less than 3 MPa, the desulfurization activity of the catalyst is lowered and the hue of the produced oil is deteriorated. Conversely, when the pressure exceeds 8 MPa, the equipment cost and the operating cost are increased to increase the pressure resistance of the equipment. It becomes bulky and uneconomical.
When the reaction temperature is less than 220 ° C, the desulfurization activity of the catalyst is low. When the reaction temperature exceeds 380 ° C, the desulfurization activity is saturated. In addition, the equipment cost and operation cost increase.
Liquid space velocity is 5.0 hr-1If the ratio exceeds 1, the contact time between the catalyst and the raw material oil becomes too short, and the desulfurization reaction is not sufficiently performed.-1If it is less than this, the contact time becomes excessively longer than necessary, and the processing efficiency decreases.
If the hydrogen / oil ratio is less than 100 L / L, the desulfurization reaction does not proceed sufficiently. Conversely, if it exceeds 400 L / L, excessive hydrogen is consumed, which increases the processing cost and is uneconomical. .
[0014]
As the hydrodesulfurization catalyst used in the method of the present invention, various supports can be used. For example, as a carrier, silica, alumina, boria, magnesia, titania, silica-alumina, silica-magnesia, silica-zirconia, silica-tria, silica-beryllia, silica-titania, silica-boria, alumina-zirconia, alumina-titania Inorganic oxides such as alumina-boria, alumina-chromia, titania-zirconia, silica-alumina-tria, silica-alumina-zirconia, silica-alumina-magnesia, silica-magnesia-zirconia, or a mixture of one or more of these Is mentioned.
Among these inorganic oxides, preferable examples include alumina, silica-alumina, alumina-boria, and alumina-zirconia. Particularly preferable examples include alumina, silica-alumina, alumina-titania, aluminum-boria, Alumina-zirconia is exemplified, and γ-alumina is particularly preferred.
These inorganic oxides may be used alone or in combination of two or more.
[0015]
In the hydrodesulfurization catalyst used in the method of the present invention, the group 6 metal to be supported on the support is molybdenum contained in a content of 10 to 25% by mass in terms of catalyst and in terms of oxide, or in terms of catalyst and in terms of oxide. Or any of tungsten contained in a content in the range of 0.1 to 5% by mass.
If the molybdenum content is less than 10% by mass, the absolute amount of molybdenum acting as an active site is small, so that desulfurization activity does not appear. Conversely, if the content of molybdenum to be supported is more than 25% by mass, Aggregation occurs and the number of active sites decreases, and as a result, the desulfurization activity decreases.
If the tungsten content is less than 0.1% by mass, the nuclear hydrogenation activity of the hardly desulphurizable substance is not expressed. Conversely, if it is more than 5% by mass, the amount of tungsten supported is too much, and metal agglomeration occurs. Occurs, and as a result, the desulfurization activity decreases.
[0016]
Further, the Group 8 metal supported on the carrier by the hydrodesulfurization catalyst used in the method of the present invention is at least one of cobalt and nickel contained in a content range of 1 to 10% by mass in terms of catalyst and in terms of oxide. It is.
When the content of the group 8 metal is less than 1% by mass, sufficient desulfurization activity cannot be obtained, and when it exceeds 10% by mass, the desulfurization activity is saturated and no further effect of addition can be found.
Furthermore, phosphorus, boron, zinc, zirconia, etc. can be included in addition to the active metal composed of Group 6 metal and Group 8 metal as necessary.
[0017]
The average pore diameter of the finished catalyst is preferably 60 to 90 mm. When the average pore diameter is too large, the diffusibility of the sulfur compound into the pores is good, but the effective surface area of the catalyst becomes small, so that the desulfurization activity decreases.
[0018]
In the method of the present invention, the light oil fraction that has undergone the hydrodesulfurization step is further hydrotreated in the hydrotreating step. The reaction conditions of the hydrotreating step are as follows: pressure (hydrogen partial pressure) is in the range of 3-8 MPa, preferably 4-7 MPa, reaction temperature is in the range of 200-360 ° C., preferably in the range of 250-350 ° C., liquid Space velocity is 1.0-5.0hr-1Range, preferably 1.0-3.0 hr-1The hydrogen / oil ratio is in the range of 100 to 400 L / L, preferably in the range of 200 to 300 L / L.
[0019]
When the pressure (hydrogen partial pressure) is less than 3 MPa, not only the hydrogenation reaction is not sufficiently performed, but also the cetane number is not improved and the hue deteriorates. To increase the cost, the equipment cost and the operating cost increase, and it becomes uneconomical.
When the reaction temperature is less than 200 ° C., the hydrogenation activity of the catalyst is low, and when it exceeds 360 ° C., the hydrogenation activity is saturated. However, the hue of the plant becomes worse, and the equipment and operating costs increase.
Liquid space velocity is 5.0 hr-1If it exceeds 1, the contact time of the catalyst and the raw material oil becomes too short, the hydrogenation reaction is not sufficiently performed, and 1.0 hr.-1If it is less than this, the contact time becomes excessively longer than necessary, and the processing efficiency decreases.
When the hydrogen / oil ratio is less than 100 L / L, a sufficient hydrogenation reaction cannot be obtained. Conversely, when the hydrogen / oil ratio exceeds 400, excessive hydrogen is consumed, which is uneconomical because the processing cost increases.
[0020]
The hydrotreating catalyst used in the method of the present invention can be variously used as a carrier. For example, as a carrier, silica, alumina, boria, magnesia, titania, silica-alumina, silica-magnesia, silica-zirconia, silica-tria , Silica-beryllia, silica-titania, silica-boria, alumina-zirconia, alumina-titania, alumina-boria, alumina-chromia, titania-zirconia, silica-alumina-tria, silica-alumina-zirconia, silica-alumina-magnesia , Inorganic oxides such as silica-magnesia-zirconia, or a mixture of two or more thereof.
Among these inorganic oxides, preferable examples include alumina, silica-alumina, alumina-boria, and alumina-zirconia. Particularly preferable examples include alumina, silica-alumina, alumina-titania, aluminum-boria, Alumina-zirconia is exemplified, and γ-alumina is particularly preferred.
These inorganic oxides may be used alone or in combination of two or more.
[0021]
Platinum, rhodium, ruthenium, etc. can be used as the hydrogenation active metal to be contained in the hydrotreating catalyst used in the method of the present invention, and one or more of these may be used in combination.
The content of these hydrogenation active metals is preferably 0.1 to 3% by mass. If the amount is less than 0.1% by mass, the hydrogenation activity is not sufficient. Conversely, if the amount exceeds 3% by mass, improvement in the hydrogenation activity cannot be expected even if the loading is increased.
Further, nickel and tungsten can be used in place of the active metal. In this case, the nickel content is preferably 1 to 10% by mass, and the tungsten content is preferably 0.1 to 5% by mass.
If necessary, in addition to these active metals, phosphorus, boron, zinc, zirconia and the like can be included.
[0022]
As described above, the method for improving the quality of cracked gas oil according to the present invention includes hydrodesulfurization of a gas oil having a specific boiling range, for example, a cracked gas oil fraction under specific conditions, and then further hydrogenation under specific conditions. This is a method for improving the quality of a gas oil fraction having a high cetane number and a good color and having a low residual sulfur content. In addition, it is preferable to provide a known hydrogen sulfide removing device between the hydrodesulfurization step and the hydrotreatment step to remove hydrogen sulfide generated in the hydrodesulfurization step.
In addition, the method of the present invention can use a normal reaction apparatus in which a fixed bed catalyst is formed in a high-pressure flow reactor and can treat a light oil fraction under relatively mild reaction conditions. Therefore, according to the method of the present invention, for example, surplus cracked gas oil can be economically converted into a gas oil fraction such as diesel fuel.
[0023]
DETAILED DESCRIPTION OF THE INVENTION
Hereinafter, embodiments of the method of the present invention will be described specifically and in detail with reference to the accompanying drawings.
FIG. 1 is a schematic diagram showing the configuration of an example of a reaction apparatus for carrying out the method of the present invention. A
A known hydrogen
[0024]
A mixed fluid of a raw oil consisting of a gas oil fraction having a boiling point range of 160 to 340 ° C. and a hydrogen-containing gas is heated to a reaction temperature by a heating furnace, a heat exchanger or the like (not shown) and reacted from the upper part of the
[0025]
Hydrodesulfurization processPassedThe gas oil fraction is further injected with a hydrogen-containing gas as necessary, and is heated or cooled to the reaction temperature and introduced into the
[0026]
【Example】
Example 1
In order to evaluate the method of the present invention, a cracked gas oil obtained from a catalytic cracker having a boiling point range of 160 to 320 ° C. having the following properties is used as a raw material, and has the same configuration as the high-
[0027]
Properties of raw oil (decomposed light oil having a boiling range of 160-320 ° C)
Specific gravity (15/4 ° C.): 0.9046
Sulfur content: 0.05% by mass
Nitrogen content: 0.021% by mass
Distillation properties
Boiling point (IBP): 160 ° C
Boiling point (50% point): 254.5 ° C
Boiling point (FBP): 320 ° C
Viscosity (@ 30 ° C): 2.641 mm2/ S
composition
Saturated hydrocarbon: 28% by volume
1-ring aromatic compound: 32% by volume
Bicyclic aromatic compound: 38% by volume
3 or more aromatic compounds: 2% by volume
Difficult to desulfurize compounds
(4,6-dimethyldibenzothiophene): 0.002% by mass
[0028]
First, Co-Mo-W / Al as a hydrodesulfurization catalyst2O3Catalyst (CoO: MoO3: WO3= 5: 25: 3.8 (mass%)) and Pt / Al as the hydrotreating catalyst2O3A catalyst (Pt = 1% by mass) is charged into the
[0029]
Pretreatment conditions for hydrodesulfurization catalyst
Pressure (hydrogen partial pressure): 3.5 MPa
Atmosphere: Hydrogen sulfide / hydrogen mixed gas flow
Temperature: Step temperature rise
2 hours at 100 ° C
2 hours at 250 ° C
2 hours at 350 ° C
[0030]
Pretreatment conditions for hydrotreating catalyst
Pressure (hydrogen partial pressure): 3.5 MPa
Atmosphere: under hydrogen gas flow
Temperature: Step temperature rise
2 hours at 100 ° C
2 hours at 250 ° C
2 hours at 350 ° C
[0031]
Subsequently, the mixed fluid of the raw material oil heated to reaction temperature and hydrogen containing gas was introduce | transduced into the
[0032]
Reaction conditions for hydrodesulfurization process
Temperature: 300 ° C
Pressure (hydrogen partial pressure): 5 MPa
Liquid space velocity: 1.5 hr-1
Hydrogen / oil ratio: 200L / L
[0033]
Reaction conditions for hydrotreating process
Temperature: 300 ° C
Pressure (hydrogen partial pressure): 5 MPa
Liquid space velocity: 1.5 hr-1
Hydrogen / oil ratio: 200L / L
[0034]
The composition of the product oil that has undergone the hydrodesulfurization step and the composition and properties of the product oil that has undergone the hydrotreatment step are as shown in Table 1 and Table 2, respectively. In Table 2, the colors are displayed in Saybolt color.
In Table 2, it can be evaluated that the lower the sulfur content, the higher the value of cetane number, and the larger the numerical value of color, the more preferable as a blend base material for diesel fuel.
[Table 1]
[Table 2]
[0035]
Example 2
The hydrodesulfurization step and the hydrotreating step were performed under the same conditions as in Example 1 except that the cracked gas oil having the following boiling point range of 160 to 330 ° C. obtained from the catalytic cracking apparatus was used as a raw material.
The composition and properties of the product oil that has undergone the hydrodesulfurization process and the product oil that has undergone the hydrotreatment process are as shown in Tables 1 and 2, respectively. In addition, the composition and property of the product oil which passed through the hydrodesulfurization process of the following Example 3 to Example 12 and Comparative Example 1 to Comparative Example 4 and the hydrotreating process were the same as in Example 2. They are shown in Table 1 and Table 2, respectively.
[0036]
Properties of raw oil (decomposed light oil having a boiling range of 160 to 330 ° C)
Specific gravity (15/4 ° C.): 0.9095
Sulfur content: 0.09% by mass
Nitrogen content: 0.041% by mass
Distillation properties
Boiling point (IBP): 160 ° C
Boiling point (50% point): 263 ° C
Boiling point (FBP): 330 ° C
Viscosity (@ 30 ° C.): 2.821 mm2/ S
composition
Saturated hydrocarbon: 28% by volume
1-ring aromatic compound: 30% by volume
Bicyclic aromatic compound: 38% by volume
3 or more aromatic compounds: 4% by volume
Difficult to desulfurize compounds
(4,6-dimethyldibenzothiophene): 0.01% by mass
[0037]
Example 3
The hydrodesulfurization step and the hydrotreating step were carried out under the same conditions as in Example 1 except that the cracked light oil having the following boiling point range of 160 to 340 ° C. obtained from the catalytic cracking apparatus was used as a raw material.
[0038]
Properties of raw material oil (decomposed light oil with a boiling range of 160-340 ° C)
Specific gravity (15/4 ° C.): 0.9143
Sulfur content: 0.13 mass%
Nitrogen content: 0.053 mass%
Distillation properties
Boiling point (IBP): 160 ° C
Boiling point (50% point): 270 ° C
Boiling point (FBP): 340 ° C
Viscosity (@ 30 ° C): 3.105mm2/ S
composition
Saturated hydrocarbon: 28% by volume
1-ring aromatic compound: 26% by volume
Bicyclic aromatic compound: 39% by volume
3 or more aromatic compounds: 7% by volume
Difficult to desulfurize compounds
(4,6-dimethyldibenzothiophene): 0.015% by mass
[0039]
Example 4
The hydrodesulfurization step and the hydrotreating step were performed on the same cracked diesel oil as in Example 1 under the same conditions as in Example 1 except that the reaction temperature in the hydrodesulfurization step was changed to 250 ° C instead of 300 ° C in Example 1. Carried out.
Example 5
In the hydrocracking gas oil as in Example 1, except that the pressure (hydrogen partial pressure) and the hydrogen / oil ratio in the hydrodesulfurization step were changed to 6 MPa and 100 L / L, respectively, instead of 5 MPa and 200 L / L in Example 1. The hydrodesulfurization step and the hydrotreatment step were performed under the same conditions as in Example 1.
Example 6
The reaction temperature and liquid space velocity of the hydrodesulfurization step were set to 300 ° C. and 1.5 hr of Example 1.-1Instead of 350 ° C. and 4.0 hr, respectively.-1The hydrodesulfurization step and the hydrotreating step were performed on the same cracked diesel oil as in Example 1 under the same conditions as in Example 1 except that.
[0040]
Example 7
The same hydrocracked gas oil as in Example 2 was used in Example 2 except that the pressure (hydrogen partial pressure) in the hydrodesulfurization step was changed to 5 MPa instead of 5 MPa in Example 2 (that is, the same pressure as in Example 1). The hydrodesulfurization step and the hydrotreatment step were performed under the same conditions (that is, the same conditions as in Example 1).
[0041]
Example 8
The same hydrodesulfurization process as in Example 3 except that the pressure in the hydrodesulfurization step (hydrogen partial pressure) was set to 7 MPa instead of 5 MPa in Example 3 (that is, the same pressure as in Example 1). The hydrodesulfurization step and the hydrotreatment step were performed under the same conditions (that is, the same conditions as in Example 1).
[0042]
Example 9
The catalyst used in the hydrodesulfurization step was Co-Mo-W / Al of Example 1.2O3(CoO: MoO3: WO3= 5: 25: 3.8 (mass%)) instead of Co-Mo / Al2O3(CoO: MoO3= 5: 20 (mass%)) The hydrodesulfurization step and the hydrotreating step were performed on the same cracked diesel oil as in Example 1 under the same conditions as in Example 1.
[0043]
Example 10
The catalyst used in the hydrodesulfurization step was Co-Mo-W / Al of Example 2.2O3(CoO: MoO3: WO3= 5: 25: 3.8 (mass%)) (ie, the same catalyst as in Example 1)2O3(CoO: MoO3= 5: 20 (mass%)) The hydrodesulfurization step and the hydrotreating step were performed on the same cracked diesel oil as in Example 2 under the same conditions as in Example 2 (that is, the same conditions as in Example 1). Carried out.
[0044]
Example 11
The catalyst used in the hydrodesulfurization step was Co-Mo-W / Al of Example 3.2O3(CoO: MoO3: WO3= 5: 25: 3.8 (mass%)) (ie, the same catalyst as in Example 1)2O3(CoO: MoO3= 5: 20 (mass%)) The hydrodesulfurization step and the hydrotreating step were performed on the same cracked diesel oil as in Example 3 under the same conditions as in Example 3 (that is, the same conditions as in Example 1). Carried out.
[0045]
Comparative Example 5
The catalyst used in the hydrotreating process was the Pt / Al of Example 1.2O3Ni-W / Al instead of (Pt = 1% by mass)2O3(NiO: WO3= 5: 20 (mass%)), and the same hydrocracking gas oil as in Example 1 except that the reaction temperature in the hydrotreating process was changed to 360 ° C. instead of 300 ° C. in Example 1. The hydrodesulfurization process and the hydrotreatment process were carried out under the conditions.
[0046]
Example 13
The hydrodesulfurization step and hydrogenation were performed under the same conditions as in Example 1 on the same cracked gas oil as in Example 1, except that the pressure (hydrogen partial pressure) in the hydroprocessing step was changed to 3 MPa instead of 5 MPa in Example 1. Processing steps were performed.
[0047]
Comparative Example 1
The catalyst used in the hydrodesulfurization treatment was obtained from the cracked gas oil obtained from the catalytic cracking apparatus and having the following boiling point range of 160 to 360 ° C. as the raw material, and the Co—Mo—W / Al of Example 1 was used.2O3(CoO: MoO3: WO3= 5: 25: 3.8 (mass%)) instead of Co-Mo / Al2O3(CoO: MoO3= 3: 15 (mass%)) The same method as in Example 1 was performed.
[0048]
Properties of raw oil (decomposed light oil with a boiling range of 160-360 ° C)
Specific gravity (15/4 ° C.): 0.9241
Sulfur content: 0.18% by mass
Nitrogen content: 0.059% by mass
Distillation properties
Boiling point (IBP): 160 ° C
Boiling point (50% point): 281.5 ° C
Boiling point (FBP): 360 ° C
Viscosity (@ 30 ° C): 3.908mm2/ S
composition
Saturated hydrocarbon: 28% by volume
1-ring aromatic compound: 24% by volume
Bicyclic aromatic compound: 37% by volume
3 or more aromatic compounds: 11% by volume
Difficult to desulfurize compounds
(4,6-dimethyldibenzothiophene): 0.025% by mass
[0049]
Comparative Example 2
The catalyst used in the hydrodesulfurization process using a cracked gas oil having a boiling point range of 140 to 300 ° C. having the following properties as a raw material was used as the Co—Mo—W / Al of Example 1.2O3(CoO: MoO3: WO3= 5: 25: 3.8 (mass%)) instead of Co-Mo / Al2O3(CoO: MoO3= 3: 15 (mass%)) The hydrodesulfurization step and the hydrotreatment step were performed under the same conditions as in Example 1.
[0050]
Properties of raw oil (decomposed light oil having a boiling range of 140-300 ° C)
Specific gravity (15/4 ° C): 0.9003
Sulfur content: 0.04 mass%
Nitrogen content: 0.02% by mass
Distillation properties
Boiling point (IBP): 140 ° C
Boiling point (50% point): 236 ° C
Boiling point (FBP): 300 ° C
Viscosity (@ 30 ° C): 2.498mm2/ S
composition
Saturated hydrocarbon: 30% by volume
1-ring aromatic compound: 34% by volume
Bicyclic aromatic compound: 35% by volume
3 or more aromatic compounds: 1% by volume
Difficult to desulfurize compounds
(4,6-dimethyldibenzothiophene): 0.001% by mass
[0051]
Comparative Example 3
The catalyst used in the hydrodesulfurization treatment using the cracked gas oil having the same boiling point range of 160 to 340 ° C. as in Example 3 as the raw material was used as the Co—Mo—W / Al of Example 1.2O3(CoO: MoO3: WO3= 5: 25: 3.8 (mass%)) instead of Co-Mo / Al2O3(CoO: MoO3= 3: 15 (mass%)), and the hydrodesulfurization treatment and hydrogenation were performed under the same conditions as in Example 1, except that the reaction temperature of hydrodesulfurization was changed to 390 ° C instead of 300 ° C in Example 1. Processing steps were performed.
[0052]
Comparative Example 4
The hydrodesulfurization step and the hydrotreating step were carried out under the same conditions as in Example 1 except that the cracked gas oil having a boiling range of 160 to 360 ° C. obtained from the catalytic cracking apparatus was used as the raw material.
[0053]
As shown in Table 2, Example 1 in which a predetermined cracked gas oil was treated under the conditions defined by the method of the present inventionTo 11 andIn Example 13, the cracked gas oil can be converted into a gas oil fraction having a residual sulfur content of 0.4% by mass or less, a high cetane number, and excellent color.
[0054]
On the other hand, in Table 2, Comparative Example 1 having a boiling range of 160 to 360 ° C. and an upper limit higher than the upper limit of the boiling range of 160 to 340 ° C. defined in the present invention is a raw material oil within the boiling range defined by the method of the present invention. Examples 1 to11 andCompared to 13, the cetane number and color were inferior to wrinkles, and the aromatic compound content of Comparative Example 111 andCompared to 13, it is much more.
Moreover, the comparative example 2 whose boiling point range is 140-360 degreeC and whose minimum is lower than the minimum of the boiling point range prescribed | regulated by this invention is Example 1-.11 and 13Compared with, the cetane number and color are inferior to wrinkles.
Comparative Example 3 in which the reaction temperature of the hydrodesulfurization step exceeds 390 ° C. and exceeds the range of the reaction temperature defined in the present invention is a reaction temperature range 220 to 380 of the hydrodesulfurization step defined by the method of the present invention. Treated Example 111 andCompared to 13, the color is particularly inferior.
Further, Comparative Example 4 having a boiling point of 160 to 360 ° C. having an upper limit higher than the upper limit of the boiling point range defined in the present invention is described in Examples 1 to 4.11 andCompared to 13, the color is remarkably inferior, and the content of the aromatic compound of Comparative Example 1 is the same as in Examples 1 to11 andCompared to 13, it is much more.
In Comparative Example 5, the reaction temperature of the hydrotreating process is changed to 360 ° C. instead of 300 ° C. in Example 1. As a result, in order to exceed the reaction temperature range of 200 to 350 ° C. defined in the present invention, the cetane number is lower than those in Examples 1 to 11 and 13, and the monocyclic aromatic content is high, and the saturated volume. % Is low.
[0055]
【The invention's effect】
According to the method for improving the quality of cracked gas oil according to the present invention, a special separator and reactor are not required, and a gas oil having a specific boiling range can be obtained under relatively mild processing conditions using an ordinary reactor. By treating under specific conditions, a high-quality gas oil fraction having a high cetane number, an excellent hue, and a low sulfur content can be obtained.
This makes it possible to economically produce a shortage of diesel fuel by using surplus cracked light oil.
[Brief description of the drawings]
FIG. 1 is a schematic diagram showing the configuration of a high-pressure flow reactor.
[Explanation of symbols]
10 High-pressure flow reactor
12, 14 reactor
16, 18 Fixed bed type catalyst layer
20 Hydrogen sulfide removal equipment
22 Gas-liquid separator
Claims (3)
白金、ロジウム、ルテニウムから選ばれた少なくとも1種類以上の金属を含んでなる水素化触媒、又はニッケル及びタングステンを含んでなる水素化触媒の存在下で3〜8MPaの範囲の水素分圧、200〜350℃の範囲の反応温度、1.0〜5.0hr−1の範囲の液空間速度、100〜400L/Lの範囲の水素/オイル比の反応条件で前記水素化脱硫工程を経た軽油留分を水素化処理して、セタン価が45〜53、セーボルト色値が2〜12、かつ1環芳香族含有量が13.0〜29.0容積%、飽和分含有量が69.5〜86.5容積%、2環芳香族含有量が0.5〜2.2容積%の生成油を生成するようにした水素化処理工程と
を備えていることを特徴とする軽油の高品質化処理法。Reaction with hydrogen partial pressure in the range of 3-8 MPa, reaction temperature in the range of 220-380 ° C., liquid space velocity in the range of 1.0-5.0 hr −1 and hydrogen / oil ratio in the range of 100-400 L / L. A desulfurization catalyst in which a group 6 metal and a group 8 metal are supported on a carrier made of an inorganic oxide under the conditions, and a boiling point range of 160 to 340 ° C, and an upper limit of the boiling point range is 320 ° C or higher and 340 ° C or lower A desulfurization step in which a light oil fraction is hydrodesulfurized by contact reaction with the light oil fraction;
Hydrogen partial pressure in the range of 3-8 MPa in the presence of a hydrogenation catalyst comprising at least one metal selected from platinum, rhodium, ruthenium, or a hydrogenation catalyst comprising nickel and tungsten, 200- A gas oil fraction that has undergone the hydrodesulfurization step under reaction conditions of a reaction temperature in the range of 350 ° C., a liquid space velocity in the range of 1.0 to 5.0 hr −1 , and a hydrogen / oil ratio in the range of 100 to 400 L / L. The cetane number is 45 to 53 , the Saebold color value is 2 to 12 , the monocyclic aromatic content is 13.0 to 29.0 % by volume, and the saturated content is 69.5 to 86. And a hydrotreating step for producing a product oil having a content of 5 % by volume and a bicyclic aromatic content of 0.5 to 2.2 % by volume. Law.
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JP35160795A JP3634041B2 (en) | 1995-12-26 | 1995-12-26 | Light oil quality treatment method |
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JP35160795A JP3634041B2 (en) | 1995-12-26 | 1995-12-26 | Light oil quality treatment method |
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FR2778341B1 (en) * | 1998-05-07 | 2000-06-09 | Inst Francais Du Petrole | CATALYST BASED ON NOBLE GROUP VIII METAL CONTAINING BORON AND / OR SILICON AND ITS USE IN HYDROCARBON CHARGE HYDROTREATMENT |
US6162956A (en) * | 1998-08-18 | 2000-12-19 | Exxon Research And Engineering Co | Stability Fischer-Tropsch diesel fuel and a process for its production |
WO2001081506A1 (en) * | 2000-04-20 | 2001-11-01 | Exxonmobil Research And Engineering Company | Production of low sulfur distillates |
US8632675B2 (en) * | 2008-12-24 | 2014-01-21 | Exxonmobil Research And Engineering Company | Co-processing of diesel biofeed and heavy oil |
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