脱丙烯精馏塔

1.设计题目：试设计一座分离乙烷和丙烯的板式连续精馏塔。2.设计任务物料处理量10万吨/年进料组成组分CH4C2H6C3H6C3H8C4H10总合组成0.050.350.150.200.251.00分离要求:塔顶产品:丙烯含量2%不出现丙烷及更重组分塔底残液:乙烷含量2%不出现甲烷塔操作条件:平均操作压力:27.4atm进料热状况:饱和液体进料进料温度:26℃回流比:自选单板压降:≦0.7kPa塔板类型:自选工作日:每年300天,每天24小时连续运行3.1.2清晰分隔物料衡算确定轻重关键组分,选取C2H6为轻关键组分,C3H6为重关键组分。由于精馏的任务是把C2H6、C3H6与CH4、C3H8、C4H10混合物分开,按清晰分割情况确定各组分在塔顶、进料和塔底的数量，组成以及操作温度。3.1.3计算塔顶塔底组成,塔顶塔底温度1.各组分平均摩尔质量0.0516.040.3530.700.1542.0810.2044.0970.2558.12440.99/MkgKmol进料量F=81.010(30024)338.84/MKmolh由进料组成，进料量按清晰分割求'D，'W1.F=338.84Kmol/h0.02DhX0.02WlX2.乙烷为轻关键组分，丙烯为重关键组分。3.338.840.35118.594/338.840.1550.826/lhfKmolhfKmolh4.计算1Liifhiihf1338.84(0.050.35)135.54Liif203.30hiihf5.0.02135.54203.3010.0210.020.020.0210.024.094LW118.5944.094114.5llldfW0.02((4.094135.54)110.022.68DhhLlDhXdWfX50.8262.6848.146hhhWfd114.527.974.0942.680.05648.146lLhhdWdWD=114.5+2.68+338.84×0.05=134.122W=4.094+48.146+338.84×(0.25+0.20)=204.721.塔顶温度Dt。由露点方程计算查2.74MpaT=397.4siap设1℃组分CH4C2H6C3H6C3H8C4H10DiX0.12630.85370.0199800组分CH4C2H6C3H6C3H8C4H10iZ0.050.350.150.200.251.00if16.942118.59450.82667.76884.71338.84iD16.942114.52.6800134.122DiX0.12630.85370.01998001iW04.09448.14667.76884.71204.72wiX00.0199980.2350.3310.41381iK50.940.30.260.071.00051.01iiyk2.塔底温度wt.由泡点方程：wt=82℃=179.6°F组分CH4C2H6C3H6C3H8C4H10wiX00.0199980.2350.3310.4138iK8.62.81.351.250.510.998ix不清晰分割验证求以重关键组分为对比组分的各组分的平均相对挥发度计算列表如下：iDiKihDaWiKihWaihaCH4516.678.66.3710.30C2H60.943.132.82.072.55C3H60.311.3511C3H80.260.871.250.930.81C4H100.070.230.510.380.30iDihDhKaKiWihWhKak1/2(.)ihihdihWaaa代入汉斯特别克公式，得到loglog(/)log(/)log(/)log(/)logloglg0.056lg27.97log0.056log2.55log1.251.4471.250.4071.256.63logiHiHlHLHihihihadwdwdwdwaaaa以重关键组分丙烷为对比组分,分别将除关键组分以外的各组分的平均相对挥发度iha代入上式求得(/)idw进一步求得idiwDixwix列表如下：ifi(/)idwidiwDixwixiha16.942CH42.92×51016.93255.8100.12546.141010.30118.594C2H628.18114.534.0640.850.01992.5550.826C3H60.0562.69548.1250.01990.236167.768C3H80.01380.92266.8460.00680.3280.8184.71C4H1051.911031.621084.70851.2100.4160.30338.84/135.11203.741.0001.000/0.019DHX(小于2％）0.019WLX(小于2％）均小于规定的浓度值符合要求。3.1.4由恩德伍德方程计算Rmin塔顶塔底平均温度是：T=50℃.以重组分C3H8为对比组分，求各组分的相对挥发度查各组分在397.4MPa50℃下的K值列表计算如下:iiZiKihaCH40.057.48.6C2H60.3522.33C3H60.150.861C3H80.200.770.9C4H100.250.280.33由于是泡点进料所以e=0由1ciiiiaZeae=0,通过试差计算求θ列表计算组分iZiiiaZa1.351.3561.36CH40.058.60.0590.0590.055C2H60.352.330.8370.8420.845C3H60.1510.4290.4210.417C3H80.200.90.40.3950.39C4H100.250.330.080.0810.081计算得=1.356各组分塔顶含量如下表iCH4C2H6C3H6C3H8C4H10dix0.1250.850.01990.006801ia30.437.833.172.911112.07ciDimiiaxRa1.07mR所以该塔最小回流比为1.073.3由芬斯克方程计算mN0.85dlx0.0199wlx0.0199dhx0.236whx求塔顶。塔底温度，压力为2.74Mp条件下的相对挥发度，计算列表如下：温度iKiKlhaC2H6C3H61dt0.940.33.13w82t2.81.352.07,,2.25lhlhDlhWaaalog(0.86/0.014)(0.24/0.008)log2.256.657mN块所以最小理论板数为8块3.3由经验公式确定理论塔板数操作回流比一般定为最小回流比的1.2---2倍，取R=1.6mR=1.81.81.10.2512.8mRRxR0.391TmTNNyN查吉利兰图得y=0.3912.1TN块3.4由奥康奈尔图确定板效率该塔平均操作温度1/2()50dwttt列表计算当P=2.74Mpa50℃查得2lk0.86hkiFiixZli.FilixCH40.0500C2H60.3500C3H60.150.0780.0117C3H80.200.0860.0172C4H100.250.1470.03681.000/0.066则2.332.330.0660.154lhlhlaa由奥康奈尔图查得:总板效率0.77T3.5确定进料板位置(1)实际塔板数取进料位置扣除再沸器以后计算实际塔板数(1)/14.415aTTNN块取进料位置()()1aRmSmNNNlog(/).(/)()0.88()log(/).(/)lhdhlfRmSmlhfhlwxxxxNNxxxx12.1=()SmN+0.88()SmN+1得到()SmN=5.9块()12.15.95.2RmN块精馏段实际塔板数为7块。提馏段实际塔板数为8块。可在自下而上第7层开进料口。3.6塔工艺的计算结果精馏塔工艺计算结果一览表项目符号数值单位进料流量F406.6Kmol进料温度t26℃操作压力P2.74MPa塔顶产品流量D135.41塔顶温度1℃塔底产品流量W203.74塔底温度82℃最小回流比1.14.1.1板型选取根据化学工业出版社《化工原理》提供的液相流量参考表选取单流型塔板，单流型塔板是最常用的形式，结构简单，制作方便，且横贯全板的流道长，有利于达到较高的塔板效率。4.1.2板间距的初选板间距NT的选定很重要，对完成一定生产任务若采用较大的板间距能允许较高的空塔气速，对塔板效率、操作弹性及安装检修有利，但板间距增大后会增加塔身总高度金属消耗量，塔基、支座等的负荷，从而导致全塔造价增加。反之，采用较小的板间距只能允许较小的空塔气速，塔径就要增大，但塔高可降低。但是板间距过，小容易产生液泛现象降低板效率。所以在选取板间距时要根据各种不同情况予以考虑。如对易发泡的物系板间距应取大一些以保证塔的分离效果。板间距与塔径之间的关系，应根据实际情况结合经济权衡，反复调整已做出最佳选择。设计时通常根据塔径的大小由塔板间距的经验数值选取.初选板间距为0.50m.4.2汽、液体体积流量计算4.2.1精馏段、提馏段的摩尔流量计算精馏段气体摩尔流量V=L+D=(R+1)D=(1.8+1)×134.122=375.54kmol/h=0.104kmol/s提馏段气体摩尔流量v′=v=0.104kmol/s实际回流比R1.8最少理论板数7块全塔理论板数N12块全塔平均板效率77%精馏段实际塔板数7块提馏段实际塔板数8块全塔实际板数15块精馏段液体摩尔流量L=RD=1.8×162.64=292.75kmol/h=0.081kmol/s提馏段液体摩尔流量L′=L+F=292.75+338.84=631.59kmol/h=0.18kmol/s4.2.2精馏段、提馏段的体积流量计算表4-1气体体积流量计算表组分iCH4C2H6C3H6C3H8C4H10MiM1630424458DiX0.1250.850.01990.00680WiX00.01990.2360.3280.416DiiXM225.50.840.027028.37WiiXM00.609.914.4324.1349.06/DiiXk0.0250.900.0660.0260iWikX00.0560.320.410.21iiwiMkX01.6813.4418.04812.1845.34/iDiiMXk0.4272.7721.144031.32塔顶气体密度：31274031.3237.65/(273.151)8.314VvlpMkgmRT塔底气体密度：32274045.3442.07/(273.1582)8.314VvlpMkgmRT气体平均密度：31237.6542.0739.86/22vvvkgm塔顶气体体积流量：31110.10431.320.09/37.65vSvVMVms塔底气体积流量：32220.10445.340.11/42.07vSvVMVms全塔平均气体体积流量：3120.090.110.1/22SSSVVVms按塔底温度85℃计算液体体积流量表4-2液体体积流量计算表3311526/1.910Liikgmm塔顶液体体积流量：331110.08145.346.8910/526SLLMLms塔顶液体体积流量：31220.1931.90.01/580SLLMLms组分iCH4C2H6C3H6C3H8C4H10MWiX00.01990.2360.3280.416iM1630424458WiiXM00.609.914.4324.1349.06wiiiwiiX