《化工原理课程设计》板式精馏塔 设计报告

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《化工原理课程设计》报告

4万吨/年甲醇~水板式精馏塔设计

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目录一、概述························································································41.1设计依据··················································错误！未定义书签。1.2技术来源··················································错误！未定义书签。1.3设计任务及要求···································································5二：计算过程··················································································61.塔型选择··············································································72.操作条件的确定·····································································72.1操作压力······································································72.2进料状态······································································72.3加热方式······································································72.4热能利用······································································73.有关的工艺计算·····································································83.1最小回流比及操作回流比的确定··············错误！未定义书签。3.2塔顶产品产量、釜残液量及加热蒸汽量的计算错误！未定义书签。3.3全凝器冷凝介质的消耗量···············································153.4热能利用············································错误！未定义书签。3.5理论塔板层数的确定·····················································153.6全塔效率的估算···································错误！未定义书签。3.7实际塔板数NP····································错误！未定义书签。4.精馏塔主题尺寸的计算··································错误！未定义书签。4.1精馏段与提馏段的体积流量····················错误！未定义书签。4.1.1精馏段·······························································184.1.2提馏段·······························································204.2塔径的计算·································································224.3塔高的计算·································································305.塔板结构尺寸的确定·····························································245.1塔板尺寸····································································255.2弓形降液管·········································错误！未定义书签。5.2.1堰高··································································265.2.2降液管底隙高度h0················································275.2.3进口堰高和受液盘························错误！未定义书签。5.3浮阀数目及排列···························································28

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5.3.1浮阀数目····························································285.3.2排列··································································285.3.3校核··································································296.流体力学验算······································································306.1气体通过浮阀塔板的压力降(单板压降)hp··························306.1.1干板阻力hc························································306.1.2板上充气液层阻力h1············································316.1.3由表面张力引起的阻力h?······································316.2漏液验算····································································316.3液泛验算····································································316.4雾沫夹带验算······························································327.操作性能负荷图···································································327.1雾沫夹带上限线···························································327.2液泛线·······································································337.3液体负荷上限线···························································337.4漏液线·······································································337.5液相负荷下限线···························································347.6操作性能负荷图···························································348.各接纳尺寸的确定································································368.1进料管·······································································368.2釜残液出料管······························································368.3回流液管····································································378.4塔顶上升蒸汽管···························································378.5水蒸汽进口管······························································383

一、概述1.1设计背景塔设备是化工、炼油生产中最重要的设备之一。塔设备的设计和研究，已经受到化工行业的极大重视。在化工生产中，塔设备的性能对于整个装置的产品产量、质量、生产能力和消耗定额，以及三废处理和环境保护等各个方面，都有十分重大的影响。精馏过程的实质是利用混合物中各组分具有不同的挥发度。即在同一温度下，各组分的饱和蒸汽压不同这一性质，使液相中的轻组分转移到汽相中，汽相中的重组分转移到液相中，从而达到分开的目的。因此精馏塔操作弹性的好坏直接关系到石油化工企业的经济效益。为了加强工业技术的竞争力，长期以来，各国都在加大塔的研究力度。如今在我国常用的板式塔中主要为泡罩塔、浮阀塔、筛板塔和舌型塔等。填料种类出拉西、环鲍尔环外，阶梯环以及波纹填料、金属丝网填料等规整填料也常采用。更加强了对筛板塔的研究，提出了斜空塔和浮动喷射塔等新塔型。同时我国还进口一些新型塔设备，这些设备的引进也带动了我国自己的塔设备的科研、设计工作，加速了我国塔技术的开发。国外关于塔的研究如今已经放慢了脚步，是由于已经研究出了塔盘的效率并不取决与塔盘的结构，而是主要取决与物系的性质，如：挥发度、黏度、混合物的组分等。国外已经转向研究“在提高处理能力和简化结构的前提下，保持适当的操作弹性和压力降，并尽量提高塔盘的效率。〞在新型填料方面则在努力的研究发展有利于气液分布均匀、高效和制造便利的填料。经过我国这些年的努力，在塔研究方面与国外先进技术的差距正在不断的减小4

目前，精馏塔的设计方法以严格计算为主，也有一些简化的模型，但是严格计算法对于连续精馏塔是最常采用的，我们此次所做的计算也采用严格计算法。1.2设计条件原料：甲醇、水原料温度：泡点进料处理量：4万吨/年原料组成：甲醇的质量分率wf=0.35（质量分数）产品要求：塔顶甲醇的质量分率wd=0.94（质量分数），塔底甲醇质量分率=0.02（质量分数）生产时间：300天/年冷却水进口温度：25℃加热剂：0.9MP饱和水蒸汽单板压降：小于或等于0.7kpa生产方式：连续操作，泡点回流全塔效率：Et=50%1.3设计要求1.撰写课程设计说明书一份2.带控制点的工艺流程图一张3.塔装备的总装图一张

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1.4设计说明书的主要内容1.设计方案的确定2.带控制点的工艺流程图的确定3.操作条件的选择（包括操作压强、进料状态、加热剂、冷却剂、回流比）4.塔的工艺计算（1）全塔物料衡算（2）最正确回流比的确定（3）理论板及实际板的确定（4）塔径的计算（5）降液管及溢流堰尺寸的确定（6）浮阀数及排列方式（筛板孔径及排列方式）的确定（7）塔板滚动性能的校核（液沫夹带校核，塔板阻力校核，降液管液泛校核，液体在降液管内停留时间校核，严重漏液校核）（8）塔板负荷性能图的绘制（9）塔板设计结果汇总表5.辅助设备工艺计算（1）换热器的面积计算及选型（2）各种接纳管径的计算及选型（3）泵的扬程计算及选型6.塔设备的结构设计：（包括塔盘、裙座、进出口料管）二：计算过程6

1.塔型选择根据生产任务，若按年工作日300天，每天开动设备24小时计算，由于产品粘度较小，流量较大，为减少造价，降低生产过程中压降和塔板液面落差的影响，提高生产效率，选用浮阀塔。2.操作条件的确定2.1操作压力压力为P0?1.01325?105（Pa）2.2进料状态虽然进料方式有多种，但是饱和液体进料时进料温度不受季节、气温变化和前段工序波动的影响，塔的操作比较简单控制；此外，饱和液体进料时精馏段和提馏段的塔径一致，无论是设计计算还是实际加工制造这样的精馏塔都比较简单，为此，本次设计中采取饱和液体进料（q=1）。2.3加热方式精馏塔的设计中多在塔底加一个再沸器以采用间接蒸汽加热以保证塔内有足够的热量供应；由于乙醇~水体系中，乙醇是轻组分，水由塔底排出，且水的比热较大，故可采用直接水蒸气加热，这时只需在塔底安装一个鼓泡管，于是可省去一个再沸器，并且可以利用压力较底的蒸汽进行加热，无论是设备费用还是操作费用都可以降低。2.4热能利用精馏过程的原理是屡屡部分冷凝和屡屡部分汽化。因此热效率较低，寻常进入

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再沸器的能量只有5%左右可以被有效利用。虽然塔顶蒸汽冷凝可以放出大量热量，但是由于其位能较低，不可能直接用作为塔底的热源。为此，我们拟采用塔釜残液对原料液进行加热。3.物料的工艺计算由于精馏过程的计算均以摩尔分数为准，需先把设计要求中的质量分数转化为摩尔分数。原料液的摩尔组成：甲醇的摩尔质量为：32kg/kmol水的摩尔质量为：18kg/kmol0.35/32xF??0.230.35/32?0.65/18xD?0.94/32?0.8980.94/32?（1?0.94）/180.02/32xW??0.010.02/32?（1?0.02）/18以年工作日为300天，每天开车24小时计，进料量为：进料液的平均摩尔数MF?32?0.23?18?(1?0..23)?21.22kg/kmolm4?107F???261.8kmol/hMF21.22?24?300根据公式D?F.xF?xW0.23?0.01?261.8?xD?xW0.898?0.018

可求出D?64.86kmol/h由全塔的物料衡算方程可写出：F?D?W求得W?196.94kmol/h表1.原料液、馏出液与釜残液的流量名称w(质量分数)x(摩尔分数)原料液0.350.23261.8馏出液0.940.89864.86釜残液0.020.01196.94流量kmol/h3.1相对挥发度可根据平衡线图（图3-1）查得塔顶、塔底温度1—汽相2—液相

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图3-1甲醇-水的等压曲线或用计算法求得：①塔顶：y1?0.957，P?101.325kpao假设t=83℃,利用安托因方程lgPA?6.03055?1211.033，t?220.79lgPBo?6.07954?1344.8t?219.48oo计算得出PA?110.70kpa，PB?43.016kpaP?PBoPAoxy?再利用x?oo，PPA?PB求得x1?0.8615,y1?0.9412oo假设t=82℃，同理求得PA?107.39kpa，PB?41.58kpax1'?0.9078,y1'?0.9622利用比例差值法求出塔顶温度：t?820.957?0.9622?，则t1?82.25℃83?820.9412?0.9622当t=82.25℃时，计算得出PAo?108.21kpa，PBo?41.93kpaoPA108.21?2.581此时的相对挥发度?1?o?PB41.93②塔进料处：x2?0.54110

oo假设t=90℃,同理求得PA?136.12kpa，PB?54.233kpax2?0.5751oo假设t=91℃,同理求得PA?140.1kpa，PB?56kpa'x2?0.5389利用比例差值法求出塔进料处温度：t?900.541?0.5751?，则t2?90.94℃91?900.5389?0.5751当t=90.94℃时，计算得出PAo?139.85kpa，PBo?55.89kpaoPA139.85?2.502此时的相对挥发度?2?o?PB55.89③塔底：x3?0.035oo假设t=108℃,同理求得PA?222.46kpa，PB?93.98kpax3?0.0572oo假设t=109℃,同理求得PA?228.253kpa，PB?96.723kpa'x3?0.0350则得出塔底温度：t3?109℃当t=109℃时，oPA228.253????2.360此时的相对挥发度3oPB96.723

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全塔的相对挥发度??3.2回流比R的确定?1?2?3?2.581?2.502?2.360?2.479由于是泡点进料（q=1），xq?xF?0.541相平衡方程y??x1?(??1)x当x?xF,求出夹紧点xP?0.541，yP?0.745，因此：Rmin?xD?yP0.957?0.745??1.039yP?xP0.745?0.541操作回流比R?(1.1?2)Rmin最少理论板数Nmin的确定：利用芬斯克方程??xD??1?xWlg????1?x?D??xW??lg?????0.541??1?0.035????lg?????????1?0.541??0.035?????7.07lg2.479Nmin由于设备的综合费用与N(R+1)有直接的关系，因此绘制N(R+1)~R图就可以求当R值时N(R+1)最小的为实际R令??R，由不同β得到R值Rmin利用吉利兰图

N?NminR?Rmin~求出N值，进而能得到N(R+1)N?1R?112

吉利兰图分别取β=1.1、1.2、1.3、1.4、1.45、1.5、1.55、1.6、2，将查上图或计算出相应的值，见下表：βRR?RminR?1N?NminN?11.11.1430.0490.591.21.2470.0930.521.31.3510.1330.511.41.45460.1690.4951.451.5070.1870.4913

NN(R+1)βR18.740.01.51.560.2030.4613.935.715.835.51.551.610.2190.46514.0836.7615.4736.371.61.66240.2340.4613.9437.1314.9836.7722.0780.3380.3711.836.3214.8237.16R?RminR?1N?NminN?1NN(R+1)验算：若lgR?Rmin?0.17时，可以用下公式：R?1???0.17?N?Nmin?R?Rmin??0.9?N?1?R?1①若R=1.2，则Rmin?1.039,利用公式lgN?Nmin?R?Rmin??0.9?N?1?R?1N?Nmin??0.17?0.581，则N?18.26，求得求出?N?1?N?R?1??40.172。②若R=1.3，利用公式lgN?Nmin?R?Rmin??0.9?N?1?R?1???0.17求出?N?Nmin?0.5344，则N?16.33，求得N?R?1??37.559N?1③若R=1.4，利用公式lgN?Nmin?R?Rmin??0.9?N?1?R?1N?Nmin??0.17?0.495，求出?N?1?则N?14.98，求得N?R?1??35.95。14

计算结果说明在R=（1.15~1.35）范围内R?Rmin?0.17，但N?R?1?值确随RR?1值增大而减小，无最小值，所以根据作图找到最适回流比R=(1.56~1.61).取R=1.573.3物料平衡①精馏段操作方程：y?xR1.570.957x?D?x??0.611x?0.372R?1R?12.572.57精馏段液体的摩尔流量：L?RD?1.57?45.142?70.873kmol/h气体的摩尔流量：V??R?1?D?2.57?45.142?116.015kmol/h②提馏段操作方程：液体的摩尔流量：L'?L?qF?70.873?82.255?153.128kmol/h气体的摩尔流量：V'?V??q?1?F?116.015?0?116.015kmol/hWxL'153.12837.613?0.035y?'x?W?x??1.32x?0.011'VV116.015116.015q线方程：x?0.54115

3.5理论塔板层数的确定精馏段操作线方程：y?0.611x?0.372提馏段操作线方程：y?1.32x?0.011q线方程：xf?xF?0.541相平衡方程：y??x2.479x?1????1?x1?1.479x利用逐板法计算理论塔板层数：相平衡方程y1?xD?0.957?????x1?0.900y2?0.922???x1?0.827y3?0.877???x3?0.742y4?0.825???x4?0.655y5?0.772???x5?0.577y6?0.725???x6?515?xF?0.541（进料板）y7?0.669???x7?0.449y8?0.582???x8?0.360y9?0.464???x9?0.259y10?0.331???x10?0.166

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y11?0.208???x11?0.096y12?0.116???x12?0.050y13?0.055???x13?0.023?xW?0.003从上计算中可以得出理论塔板层数N理?13块（含塔釜）其中，第6块为进料板。N理?52%由条件知全塔效率Et?N实N理13??25（含塔釜）则可计算出实际塔板层数N实?Et0.524.基本物性数据计算根据苯~甲苯系的相平衡数据可以查得：y1?xD?0.957x1?0.902(塔顶第一块板)yF?0.719xF?0.541(加料板)xw?0.035yw?0.032(塔釜)全塔的相对平均挥发度：???1?2?3?2.581?2.502?2.360?2.479全塔的平均温度：tm?tD?tF?tW82.25?90.94?109??91.115oC3317

4.1精馏段整理精馏段的已知数据列于表3(见下页)，由表中数据可知：表3精馏段的已知数据位置进料板'xF?0.5塔顶(第一块板)y1'?xD'?0.95x1'?0.884y1?xD?0.957x1?0.9ML1?79.4MV1?78.60282.25质量分数'yF?0.684xF?0.541摩尔分数yF?0.719MLF?84.426摩尔质量/kg/kmolMVF?81.934温度/℃90.94①液相平均摩尔质量：MLF?ML184.426?79.4??81.913kg/kmol22t?t90.94?82.25?86.6oC平均温度：tm?FD?22M?在平均温度下查得?苯?807kg/m3,?甲苯?803.4kg/m3液相平均密度为：1?Lm?'xLm?苯?'1?xLm?甲苯18

其中，平均质量分数x'Lm'xF?x1'0.5?0.884???0.69222所以，?Lm?805.9kg/m3精馏段的液相负荷L?70.873kmol/hLn?LM?Lm?70.873?81.913?7.204m3/h805.9②汽相平均摩尔质量：M?MVF?MV181.934?78.602??80.268kg/kmol22压强PN?101.325?N?0.65kpa汽相平均密度为：?Vm?PMRT其中，平均压强P?P101.325?0.65?101.325?6?0.651?PF??103.6kpa22所以，?Vm?2.78kg/m3精馏段的汽相负荷V?116.015kmol/hVn?精馏段的负荷列于表4。表4精馏段的汽液相负荷名称液相汽相VM?Vm?116.015?80.268?3349.7m3/h2.7819

平均摩尔质量/kg/kmol平均密度/kg/m体积流量/m/h3381.913805.97.204(0.002m/s)380.2682.783349.7(0.93047m/s)34.2提馏段整理提馏段的已知数据列于表5，采用与精馏段一致的计算方法可以得到提馏段的负荷，结果列于表6。表5提馏段的已知数据位置塔釜'x2?0.03进料板'xF?0.5'yF?0.684质量分数'y2?0.047x2?0.035摩尔分数xF?0.541yF?0.719y2?0.055ML2?91.51摩尔质量/kg/kmolMLF?84.426MVF?81.93490.94MV2?91.23温度/℃109①液相平均摩尔质量：MLF?ML284.426?91.51??87.968kg/kmol22t?t90.94?109?95.63oC平均温度：tm?FD?22M?20

在平均温度下查得?苯?797kg/m3,?甲苯?794.4kg/m3液相平均密度为：1?Lm?'xLm?苯?'1?xLm?甲苯''xF?x20.5?0.03???0.45722其中，平均质量分数x所以，?Lm''Lm?795.6kg/m3提馏段的液相负荷L'?153.128kmol/hLn?''LM?Lm?153.128?87.968?16.931m3/h795.6②汽相平均摩尔质量：M?MVF?MV281.934?91.23??86.582kg/kmol22PMRT汽相平均密度为：?Vm?平均压强P?P2?PF101.325?13?0.65?101.325?6?0.65??107.5kpa22Vm'3??3.036kg/m所以，提馏段的汽相负荷V'?116.015kmol/hVn?'V'M?Vm?116.015?86.582?3308.6m3/h3.036表6提馏段的汽液相负荷名称

液相21

汽相

平均摩尔质量/kg/kmol平均密度/kg/m体积流量/m/h3387.968795.616.931(0.004703m/s)386.5823.0363308.6(0.9190m/s34.3全塔的流量由于精馏段和提馏段的上升蒸汽量相差不大，为便于制造，我们取两段的塔径相等。有以上的计算结果可以知道：汽塔的平均蒸汽流量：(Vn?Vn')3349.7?3308.6VS???3329.15m3/h?0.92476m3/s22汽塔的平均液相流量：(Ln?L'n)7.204?16.931LS???12.068m3/h?0.00335m3/s22汽塔的汽相平均密度：?V??Vm??'2Vm?2.78?3.036?2.908kg/m32汽塔的液相平均密度：?L??Lm??'2Lm?805.9?795.6?800.75kg/m325.塔径的计算塔径可以由下面的公式给出：22

D?4VS?u由于适合的空塔气速u?(0.6~0.8)umax，因此，需先计算出最大允许气速umax。umax?C?L??V?V初步设定板间距HT?0.45mL?12.068800.75功能参数：FLV?()L??0.06V?V3329.152.908从史密斯关联图查得：C20?0.081，由于C?C20(?20t?t?t82.25?90.94?109?91.115oC全塔的平均温度：tm?DFW?33)0.2，需先求平均表面张力：23

在此温度下，平均摩尔分数为81.913?87.968?89.9682查《化工原理》书379页液体表面张力共线图并计算出液体表面张力??20.4mN/m2史密斯关联图是按液体表面张力??20mN/m2的物系绘制的，若所处物系的表面张力为其他值，则需按式C?C20(C?0.081?(20.40.2)?0.081320?20)0.2校正查出的负荷系数，即：umax?C?L??V800.75?2.908?0.0813??1.347m/s?V2.908u=（0.6~0.8）umax=（0.808~1.077）m/s则取适合的空塔气速u?0.85m/s塔径的确定：①精馏段：气相流量Vn?3349.7m3/h?0.93047m3/s塔径D?4Vn4?0.93??1.181m?u??0.85②提馏段：气相流量Vn'?3308.6m3/h?0.9190m3/s塔径D'?根据塔径系列尺寸圆整为D?1200mm6.塔板结构尺寸的确定4Vn'?u?4?0.919?1.173m??0.8524

6.1确定塔板的流型由于塔径大于800mm，所以采用单溢流型分块式塔板。6.2塔板尺寸选取lW，而lW/D?（0.6~0.8）所以lW?（0.72~0.96），取lW?0.8m即lW/D?0.667然后根据上表弓形降液管的宽度与面积即可查出AWd?0.13，f?0.07DAT从而计算出：25

塔板总面积AT??4D2?1.1304m2弓形溢流管宽度Wd?156mm弓形降液管面积Af?0.079m2验算：液体在精馏段降液管内的停留时间??AfHTLn?0.079?0.45?17.775s?5s0.002液体在精馏段降液管内的停留时间??'AfHTL'n?0.079?0.45?7.56s?5s0.047036.3弓形降液管6.3.1堰上液流高度how本设计采用平堰，则堰上液头高how应在（6~60mm之间）。对于平堰，则堰上液头高how可用佛兰西斯公式计算：?Lh?how?2.84?10E???lW??323对于式中液流收缩系数E可用下表差得E?1.0526

则计算how?2.84?10?1.05??3?12.068???18.2mm?6mm0.8??23当平堰上液头高how?6mm时，堰上溢流会不稳定，需改为齿形堰。6.3.2堰高采用平直堰，一般应使塔板上得清夜层高度hL?50~100mm，而清夜层高度hL?hW?hOW，因此有：50?hOW?hW?100?hOW取hW?50mm，则hL?50?18.2?68.2mm6.3.3溢流管底与塔盘间距离h0因hW?h0?6mm，而hW?50mm若取精馏段取h0?30mm，那么液体通过降液管底隙时的流速为u0?

LS0.00335??0.14m/s?0.1m/s（舍弃）lWh00.7?0.0327

若取精馏段取h0?42mm，那么液体通过降液管底隙时的流速为u0?LS0.00335??0.0997m/s?0.1m/slWh00.7?0.042u0的一般经验数值为0.07~0.25m/s所以取h0?42mm6.4浮阀数目及排列采用F1型重阀，重量为33g，孔径为39mm。6.4.1浮阀数目阀孔数n取决于操作时的阀孔气速u0，而u0由阀孔动能因数F0决定。浮阀数目n?4VS?d02u0气体通过阀孔时的速度u0?F0?v一般F0?8~11，对于不同工艺条件，也可以适当调整。取动能因数F?11，那么u0n?11?6.451m/s，因此2.9083329.15?4?120个??0.0392?6.4516.4.2排列阀孔的排列方式有正三角形排列和等腰三角形排列。

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若按等边三角形排列：孔心距t?d0150mm）0.907（常用有：75mm，100mm，125mm，A0/APn?d02120?3.1415?0.0392??0.143m2阀孔面积：A0?44??x??开孔鼓泡区面积：AP?2?xr2?x2?r2sin?1????r???D??Wd?Ws??0.6??0.156?0.05??0.394m2Dr??Wc?0.6?0.025?0.575m2x???0.394??则计算可得到AP?2?0.394?0.5752?0.3942?0.5752sin?1?????0.575???取t'?80mm时画出的阀孔数目只有60个，不能满足要求，取t'?65mm画出阀孔的排布图如图1所示，其中t?75mm,t'?65mm图中，通道板上可排阀孔41个，弓形板可排阀孔24个，所以总阀孔数目为N?41?24?2?89个6.4.3校核4VS?10.38m/s2?d0N气体通过阀孔时的实际速度：u0?实际动能因数：F0?10.38?1.0335?10.55(在9~12之间)开孔率：?d02N阀孔面积??(0.039)2?89?100%??100%??13.5%塔截面积4AT4?0.7854

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开孔率在10%~14之间，满足要求。4.3塔高的计算塔的高度可以由下式计算：Z?HP?(N?2?S)HT?SHT?HF?HW已知实际塔板数为N?40块，板间距HT?0.4m由于料液较清洁，无需经常清洗，可取每隔8块板设一个人孔，则人孔的数目S为：S?40?1?4个8取人孔两板之间的间距HT?0.6m，则塔顶空间HD?1.2m，塔底空间HW?2.5m，进料板空间高度HF?0.5m，那么，全塔高度：Z?1.2?(40?2?4)?0.4?4?0.6?0.5?2.5?20.2m6.流体力学验算6.1气体通过浮阀塔板的压力降(单板压降)hp气体通过浮阀塔板的压力降(单板压降)hp?hc?h1?h?6.1.1干板阻力hc浮阀由部分全开转为全部全开时的临界速度为uoc：uoc?1.82573.1/?V1.82573.1/1.0335?10.32m/s

30

由于uoc?uo?10.38m/s2?Vu01.0335?10.382所以hc?5.34?5.34??0.0367m2?Lg2?863?9.816.1.2板上充气液层阻力h1取板上液层充气程度因数??0.5，那么：h1??hL?0.5?0.06?0.03m6.1.3由表面张力引起的阻力h?由表面张力导致的阻力一般来说都比较小，所以一般状况下可以忽略，所以：hp?0.0367?0.03?0.667m?0.667?863?9.81?564.7Pa6.2漏液验算动能因数F0?5，相应的气相最小负荷VSmin为：VSmin??4d02Nu0min其中u0min?F所以VSmin??V?5/1.0335?4.92m/s2?0.0390?89?4.92?0.523m3/s?1.103m3/s?4可见不会产生过量漏液。6.3液泛验算溢流管内的清液层高度Hd?hp?hd?hL?h?

31

其中，hp?0.0667m,hL?0.06m所以，Hd?0.667?0.06?0.003?0.1297m为防止液泛，寻常Hd??(HT?hw)，取校正系数??0.5，则有：?(HT?hw)?0.5?(0.4?0.05)?0.225m可见，Hd??(HT?hw)，即不会产生液泛。6.4雾沫夹带验算VS?V?L??V?1.36LSZL泛点率=KCFAb查得物性系数K?1.0，泛点负荷系数CF?0.097ZL?D?2Wd?1?2?0.146?0.708mAb?AT?2Af?0.7854?2?0.0706?0.6442m2所以，1.103?1.0335?1.36?0.00146?0.708863?1.0335?63.4%?80%1?0.097?0.6442泛点率=可见，雾沫夹带在允许的范围之内7.操作性能负荷图7.1雾沫夹带上限线32

取泛点率为80%代入泛点率计算式，有：VS0.8??V?L??V?1.36LSZL?VSKCFAb1.0335?1.36?0.708LS863?1.03350.097?0.6442整理可得雾沫夹带上限方程为：VS?1.444?27.8LS7.2液泛线2/3液泛线方程为aVS2?b?cL2S?dLS其中，a?1.91?105?V?LN2?1.91?105?1.0335?0.0309863?86b??HT?(??1??0)?0.5?0.4?(0.5?1?0.5)?0.05?0.15c?0.1530.153??192.422lwh00.7052?0.015212/3lwd?(1??0)E(0.667)?(1?0.5)?1.02?0.667?1?3.55320.7052/3?114.9L代入上式化简后可得：VS2?4.85?6.217L2SS7.3液体负荷上限线取??5s，那么LSmax?AfHT5?0.0706?0.4?0.00565m3/s57.4漏液线

33

取动能因数F0?5，以限定气体的最小负荷：VSmin??42d0N5?V?0.523m3/s7.5液相负荷下限线L2.84?1.02?[Smin]2/3?0.006100