CN113899161B - 一种从天然气中提取乙烷的方法 - Google Patents

一种从天然气中提取乙烷的方法 Download PDF

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CN113899161B
CN113899161B CN202111184382.4A CN202111184382A CN113899161B CN 113899161 B CN113899161 B CN 113899161B CN 202111184382 A CN202111184382 A CN 202111184382A CN 113899161 B CN113899161 B CN 113899161B
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gas
ethane
tower
demethanizer
raw material
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CN113899161A (zh
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高继峰
孙娟
王艳芳
丁锋
银永明
史世杰
宋世昌
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Sinopec Oilfield Service Corp
Sinopec Zhongyuan Petroleum Engineering Design Co Ltd
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Abstract

本发明涉及一种从天然气中提取乙烷的方法，包括以下步骤：步骤一，原料气增压；步骤二，MDEA脱碳：采用活化MDEA溶液吸收增压后的原料气中的CO2，将CO2含量脱除至50ppm以下；步骤三，脱水脱汞：将步骤二中得到的脱碳后原料气，采用三塔等压分子筛脱水脱汞工艺进行脱水和浸硫活性炭脱汞，得到含水量≤1ppm、汞含量＜0.01μg/m3的净化气；步骤四，乙烷回收：结合丙烷预冷、膨胀制冷、部分干气循环工艺的乙烷回收工艺；步骤五，凝液分馏；本发明的方法采用活化MDEA脱碳技术，CO2脱除效率高，防CO2防冻堵脱甲烷塔，C2回收率达到95％以上，C3+回收率达到99％以上，产品收率高；相比于代表性的部分干气循环工艺节约能耗12.5％，经济效益显著。

Description

一种从天然气中提取乙烷的方法
技术领域
本发明属于天然气加工工艺技术领域，具体涉及一种从天然气中提取乙烷的方法。
背景技术
天然气是一种多组分的混合气体，主要成分是烷烃，其中甲烷占绝大多数，另有少量的乙烷、丙烷和丁烷，此外一般还含有硫化氢、二氧化碳、氮和水气，以及微量的惰性气体，如氦和氩等。乙烷作为裂解制乙烯的优质原料，综合生产成本约为石脑油原料的2/3，经济效益显著。目前采用的乙烷收回技术主要有液体过冷工艺、气体过冷工艺和部分干气循环工艺。液体过冷工艺可处理含CO2较多的原料气，不需要专门设置原料气脱除CO2设施；但是乙烷收率较低，约为80％-86％。气体过冷工艺的工艺相对简单，但能耗较高，乙烷回收率较低，仅80％左右。部分干气循环工艺的乙烷收率理论计算值达到99％，但仅适用于原料气中C2含量较高、CO2含量＜1.5％的工况，其中最具代表性的是美国Ortloff公司的部分干气循环(RSV)工艺，原理是在气体过冷工艺(GSP)的基础上，将部分高压外输干气与脱甲烷塔顶气换热冷凝后，通过节流过冷进入脱甲烷塔顶提供回流，回流的外输干气可对塔顶气相进行精馏，将装置乙烷收率提升至95％以上。
某处气田伴生气中乙烷含量约225ml/m3，乙烷含量预测范围4.4-4.9％，具有较高的回收价值，但是CO2含量达到2.2％，原料中CO2含量相对较高、易造成脱甲烷塔冻堵，同时原料气气质较富会使得制冷循环能耗增加，回收率会下降。
发明内容
本发明的目的是克服现有技术的不足而提供一种适用于原料气中CO2含量大于1.5％工况的从天然气中提取乙烷的方法。
本发明的技术方案如下：
一种从天然气中提取乙烷的方法，包括以下步骤：
步骤一，原料气增压：将原料气经原料气压缩机增压至4.4MPa、48℃，得到增压原料气；
步骤二，MDEA脱碳：采用MDEA脱碳装置吸收增压后的原料气中的CO2，将CO2含量脱除至50ppm以下，得到脱碳后原料气；
步骤三，脱水脱汞：将步骤二中得到的脱碳后原料气，采用三塔等压分子筛脱水脱汞工艺进行脱水和浸硫活性炭脱汞，分子筛脱水脱汞工艺使用的分子筛脱水脱汞装置采用三个干燥塔和一个小塔，两台干燥塔在脱水时原料气上进下出，第三台干燥塔在再生时再生气下进上出，小塔用于再生气的干燥；原料气与再生气通过流量控制实现控制压差≥70kPa，再生气用量约为原料气量的10％；出塔的干气经第三粉尘过滤器过滤后，得到含水量≤1ppm、汞含量＜0.01μg/m3的净化气；
步骤四，乙烷回收：将步骤三中的净化气经主冷箱预冷至-64℃至-67℃后进入低温分进行气液分离；低温分离器分离的液相全部送至脱甲烷塔上部；低温分离器分离的气相分为两股，第一股占比14％-16％的气量经主冷箱冷至-104℃至-106℃后节流进入脱甲烷塔上部，第二股经膨胀/压缩机膨胀至2.0MPag进入脱甲烷塔上部；脱甲烷塔塔顶气相通过主冷箱复热至16℃-18℃进入膨胀/压缩机增压端增压至2.0MPag后送至外输干气增压机组；外输干气增压机组9％-12％的气量作为循环回流干气，回流干气经主冷箱冷至-104℃至-106℃送至脱甲烷塔顶部，使得回流的干气能够与脱甲烷塔内上升气相进行传质传热；脱甲烷塔塔底凝液经塔底泵增压后送出；
步骤五，凝液分馏：将脱甲烷塔底来的凝液送入脱乙烷塔中部进行分馏，脱乙烷塔塔顶气相乙烷进入乙烷冷箱冷凝后进入脱乙烷顶回流罐，然后经过脱乙烷顶回流泵增压，一部分作为乙烷产品送至乙烷储罐，另一部分回流至脱乙烷塔塔顶。
进一步的，步骤二中，MDEA溶液由重量份38.5％的MDEA、56.64％的水、4.86％的哌嗪组成；增压后的原料气送入胺液吸收塔的底部，自下而上流动的原料气与贫胺液逆流接触，将原料气中的CO2吸收脱除；胺液吸收塔顶部的净化气经第一冷却器降温、并经净化气分离器分液后进入分子筛脱水脱汞单元进行脱水。
进一步的，在脱甲烷塔的中部设置两股液相侧抽出线，两股液相侧抽出线分别经主冷箱回收冷量后返回脱甲烷塔的下一层塔盘；脱甲烷塔的下部设置一股侧抽出线，经主冷箱换热预热后回到脱甲烷塔下部。
进一步的，所述脱甲烷塔为板式塔，塔盘个数60个，三股液相侧抽出线分别从上向下数第34层塔盘、第40层塔盘和第60层塔盘处抽出。
进一步的，主冷箱连接有用于补充冷量的丙烷制冷机组，制冷温位-20℃。
进一步的，在步骤五中，乙烷产品送至乙烷储罐前，通过乙烷冷箱将乙烷产品过冷至-90℃、0.3Mpag；乙烷冷箱采用混合冷剂循环，同时利用丙烷冷剂对混合冷剂预冷，混合冷剂由体积比16％甲烷、52％丙烷和32％乙烯三组分组成，混合冷剂由混合冷剂制冷机组节流降温至-96.70℃、0.32Mpag后返回乙烷冷箱提供冷量。
进一步的，步骤五中，脱乙烷塔底设置脱乙烷重沸器，将脱乙烷塔塔底温度加热至70℃，脱乙烷塔塔底C3+凝馏送至液化气塔。
进一步的，来自脱乙烷塔的凝液首先与液化气塔底凝液通过预热器换热后进入液化气塔中部进行分馏，塔顶气进入液化气顶冷却器后去液化气顶回流罐，再经液化气顶回流泵增压后一部分作为液化气产品送至罐区，另一部分作为液化气塔顶的回流液；液化气塔塔底设置液化气重沸器，将液化气塔塔底温度加热至154℃、1.4-1.5MPag；液化气塔塔底稳定轻烃先与C3+凝液换热后再经稳定轻烃冷却器冷却至40℃后送至储罐。
进一步的，在原料气增压前后分别经入口过滤分离器和出口过滤分离器过滤和分液。
与现有技术相比，本发明的有益效果是：
本发明的方法采用活化MDEA脱碳技术，CO2脱除效率高，提高乙烷产品和天然气产品的纯度；根据原料气的组分和特性合理匹配不同温位制冷需求，节约制冷用量，减少能耗；优化部分干气循环C2+回收工艺，防CO2防冻堵脱甲烷塔，C2回收率达到95％以上，C3+回收率达到99％以上，产品收率高；本发明的方法充分考虑到原料气压力、组成、产品收率、技术指标以及技术经济等因素，投资少、能耗/操作费用低、产品收率高，相比于Ortloff公司的部分干气循环工艺节约能耗12.5％，经济效益显著。
附图说明
图1为本发明实施例的工艺流程简化示意图。
图2为本发明实施例的MDEA脱碳步骤的工艺流程图。
图3为本发明实施例的乙烷回收步骤的工艺流程图。
图4为本发明实施例的凝液分馏步骤的工艺流程图。
图5为本发明实施例的丙烷制冷机组和混合冷剂制冷机组的工艺流程图。
图中，原料气入口过滤分离器F11、原料气压缩机K1、原料气出口过滤分离器F12；
胺液吸收塔T21、空冷器AC、前净化气分离器F21、后净化气分离器F22、冷却器E23、富胺液闪蒸罐D21、贫富液换热器E21、脱碳再生塔T22、脱碳重沸器E22、再生塔回流泵P21、再生塔顶回流罐D22、胺液储罐D23、贫胺液提升泵P22、胺液循环泵P23、胺液预过滤器Q1、活性炭过滤器Q2、胺液后过滤器Q3；
干燥塔T3、第三粉尘过滤器F33、分离器三F34、预脱水冷却器AC32；
主冷箱E4、低温分离器D4、脱甲烷塔T4、膨胀/压缩机C1、塔底泵P4、第一股L1、第二股L2、回流干气的管线L3、侧抽线LC1、侧抽线LC2、侧抽出线LC3；
脱乙烷塔T5、脱乙烷重沸器E51、预热器E52、脱乙烷顶回流罐D51、脱乙烷顶回流泵P51、乙烷冷箱E5、液化气塔T6、液化气重沸器E61、液化气顶冷却器AC61、液化气顶回流罐D6、液化气顶回流泵P61、节流阀K5、稳定轻烃冷却器AC63、乙烷储罐S1、液化气储罐S2、稳定轻烃球罐S3；
外输干气增压机组J1；丙烷制冷机组J2、丙烷入口分离器J21、丙烷增压机组J22、压缩机出口冷却器J23、丙烷平衡罐J24；混合冷剂制冷机组J3、冷剂压缩机入口缓冲罐J31、混合冷剂压缩机J32、混合冷剂出口冷却器J33；BOG回收机组J4。
具体实施方式
下面将结合本发明实施例中的附图，对本发明实施例中的技术方案进行清楚、完整地描述，显然，所描述的实施例仅仅是本发明一部分实施例，而不是全部的实施例。基于本发明中的实施例，本领域普通技术人员在没有做出创造性劳动前提下所获得的所有其他实施例，都属于本发明保护的范围。
如图1至图5所示，一种从天然气中提取乙烷的方法，包括以下步骤：
步骤一，原料气增压：原料气来气压力为2.0MPa、15℃，经4台原料气压缩机K1增压增压至4.4MPa、48℃，得到增压原料气；
步骤二，MDEA脱碳：采用活化MDEA溶液吸收增压后的原料气中的CO2，将CO2含量脱除至50ppm以下，得到脱碳后原料气；
步骤三，脱水脱汞：将步骤二中得到的脱碳后原料气，采用三塔等压分子筛脱水脱汞工艺进行脱水和浸硫活性炭脱汞，分子筛脱水脱汞工艺使用的分子筛脱水脱汞装置采用三个干燥塔T3和一个小塔，两台干燥塔T3在脱水时原料气上进下出，以减少气流对床层的扰动，出塔的干气经第三粉尘过滤器F33过滤后，得到含水量≤1ppm、汞含量＜0.01μg/m3的净化气，去进行乙烷回收步骤；第三台干燥塔T3在再生时再生气下进上出，这样既可以脱除靠近干燥塔T3床层上部被吸附的物质，并使其不流过整个床层，又可以确保与湿原料气接触的下部床层得到充分再生，而下部床层的再生效果直接影响流出床层干气的露点；小塔用于再生气的干燥；原料气与再生气通过流量控制实现控制压差≥70kPa，再生气用量约为原料气量的10％，确保再生气能返回分子筛干燥器入口，减少了再生气压缩机设备的投入，同时与降压再生相比，吸收塔的抗疲劳性能强，安全性能好，使用年限长；脱碳后原料气经吸附干燥后加热至260℃，自下而上通过干燥塔T3进行再生，再生气经过空冷后进入分离器三F34，污水去排污系统，重烃去液化气塔，气相返回预脱水冷却器AC32入口；分子筛脱水工艺操作温度为35℃，操作压力4.0MPa，吸附周期16小时，再生气用量约为原料气量的10％，再生温度260℃，再生气出床层温度220℃；
步骤四，乙烷回收：将步骤三中的净化气经主冷箱E4预冷至-64℃至-67℃后进入低温分进行气液分离；低温分离器分离的液相全部送至脱甲烷塔T4上部；低温分离器分离的气相分为两股，第一股占比14％-16％的气量经主冷箱E4冷至-104℃至-106℃后节流进入脱甲烷塔T4上部，第二股经膨胀/压缩机C1膨胀至2.0MPag进入脱甲烷塔T4上部；脱甲烷塔T4塔顶气相通过主冷箱E4复热至16℃-18℃进入膨胀/压缩机C1增压端增压至2.0MPag后送至外输干气增压机组J1；外输干气增压机组J19％-12％的气量作为循环回流干气，回流干气经主冷箱E4冷至-104℃至-106℃送至脱甲烷塔T4顶部，使得回流的干气能够与脱甲烷塔T4内上升气相进行传质传热；脱甲烷塔T4塔底凝液经塔底泵P4增压后送出；
步骤五，凝液分馏：将脱甲烷塔T4底来的凝液送入脱乙烷塔T5中部进行分馏，脱乙烷塔T5塔顶气相乙烷进入乙烷冷箱冷凝后进入脱乙烷顶回流罐D51，然后经过脱乙烷顶回流泵P51增压，一部分作为乙烷产品送至乙烷储罐S1，另一部分回流至脱乙烷塔T5塔顶。
进一步的，如图1所示，在原料气压缩机K1前后设置原料气入口过滤分离器F11和原料气出口过滤分离器，进行去杂。
进一步的，如图2所示，在步骤二中，脱碳后原料气线经原料气出口过滤分离器进一步过滤分离，将粒径大于5μm的液体及杂质除去，避免杂质过多影响胺液吸收效率；之后进入MDEA脱碳装置进行脱碳处理，该工艺具有酸气负荷大、CO2脱除效率高的优点，可将原料气中CO2含量脱除至50ppm以下；MDEA溶液由重量份38.5％的MDEA、56.64％的水、4.86％的哌嗪组成，MDEA溶液简称胺液；
48.0℃、4.35MpaG的增压原料气进入胺液吸收塔T21底部，在胺液吸收塔T21内原料气自下而上流动，与贫胺液逆流接触，气体中CO2被胺液吸收脱除；胺液吸收塔T21顶部的脱碳后的原料气经空冷器AC降温，然后顺次经过前净化气分离器F21、冷却器E23冷、后净化气分离器F22分液后，去分子筛脱水脱汞装置进行脱水脱汞工艺。装置开工加入MDEA溶液时，可由胺液提升泵P22抽胺液储罐D23内配制好的溶液，送至胺液循环泵P23增压后进入胺液吸收塔T21，或由地下胺液泵将MDEA胺液储罐D23内溶液送至脱碳再生塔T22底部；
MDEA脱碳工艺还包括富胺液闪蒸、MDEA富胺液再生、酸气的冷却、装置补充水、消泡剂加入等流程，以维持MDEA脱碳装置的持续运行。具体工艺如下：
富胺液闪蒸：自胺液吸收塔T21底部来的富胺液调压至0.30MPaG后进入富胺液闪蒸罐D21，在约73℃温度下闪蒸出部分溶解在溶液中的CO2气体，闪蒸气经调压后进入尾气处理装置；
MDEA富胺液再生：经闪蒸后的富胺液进入贫富液换热器E21换热升温后进入脱碳再生塔T22顶部，与塔内自下而上的蒸汽逆流接触进行再生，解吸出溶液中的CO2气体，再生热量由设置在脱碳再生塔T22底部一侧的脱碳重沸器E22提供；脱碳再生塔T22底的热贫胺液进入贫富液换热器E21换热降温后，经过贫胺液提升泵P22输送至空冷器AC冷却至52.0℃，然后通过胺液循环泵P23增压后输送至胺液吸收塔T21顶部；胺液吸收塔T21和脱碳再生塔T22采用板式塔结构，气液接触充分，吸收效果好；贫富液换热器E21采用板式换热器，具有换热效率高，占地面积小等特点；为了对循环溶液进行过滤和保护，从脱碳再生塔T22底部出来的贫胺液经贫富液换热器E21后，经贫胺液提升泵P22和空冷后依次流经胺液预过滤器Q1、活性炭过滤器Q2及胺液后过滤器Q3，除去溶液系统中的机械杂质和降解产物，以保持溶液系统的清洁；地下胺液储罐、胺液储罐采用氮气密封，以免溶液发生氧化变质；
酸气的冷却：自脱碳再生塔T22顶部来的酸气经酸气空冷器冷却至48.0℃后进入再生塔顶回流罐D22，分离出酸性冷凝水后的酸气调压至0.04MPaG后进入尾气处理装置；分离出的酸性冷凝水由再生塔回流泵P21送至脱碳再生塔T22顶部作回流；
装置补充水：为保证胺液系统中的水平衡，需要向脱碳系统中补充0.35MPaG的脱盐水，补充量为0.93m3/h；本装置首次开工及正常生产时，配制新鲜MDEA溶液所需的脱盐水流入胺液储罐D23，再按浓度配入MDEA溶剂储存于胺液储罐D23贮备使用；
消泡剂加入：当MDEA溶液系统有严重起泡倾向或起泡时，可将消泡剂直接倒入消泡剂放大管，然后经循环后进入脱碳再生塔T22；如果消泡剂粘度较大时，可用凝结水或MDEA溶液适量稀释；消泡剂可分一次或多次注入，视溶液系统发泡情况及系统容量确定加入消泡剂量；如图2所示，消泡剂设置两处补充位置：①再生塔顶回流罐D22补脱盐水管线②贫胺液提升泵P22出口管线。
进一步的，如图1所示，由于水分与天然气在一定条件下形成水合物阻塞管路，影响冷却液化过程、也会造成不必要的动力消耗，因此需要出去天然气中的水份，步骤三中，采用分子筛脱水工艺，主要由吸附和再生两部分组成，干燥塔T3内采用浸硫活性炭，浸硫活性炭脱汞是一种经济的脱汞方法，其原理是汞与硫在反应器中反应，活性硫将汞以硫化物的方式固定在活性炭的多孔结构上，达到脱汞目的；本步骤采用湿原料气冷吹工艺，且冷吹方向与吸附方向一致，避免在冷却时床层达到一定程度的饱和对水露点的影响，达到减少分子筛吸附量及干燥器的设备尺寸的目的；同时冷吹后的高温气自下而上通过小塔进行再生。
进一步的，如图1、图3所示，步骤四乙烷回收的基本原理是低温分离，利用在一定压力下天然气中各组分的挥发度不同，将天然气制冷至露点温度以下，得到C2+凝液(NGL)，使其与甲烷等轻组分分离的过程，分离出来的NGL送至下游凝液分馏单元得到所需的产品；本单元采用了膨胀制冷+丙烷制冷，同时为了提高乙烷收率，采用了部分干气回流乙烷回收工艺技术；
具体的，来自上游的净化原料气在压力4.1MPag左右、温度35℃左右的条件下进入，净化气经主冷箱E4预冷至-65.5℃后进入低温分离器进行气液分离，液相送至脱甲烷塔T4上部；低温分离器出口气相分为两股：第一股L1约15％的气量通过主冷箱E4冷至-105℃节流进入脱甲烷塔T4上部，第二股L2通过膨胀/压缩机C1膨胀至2.0MPag、-93.6℃左右进入脱甲烷塔T4上部；脱甲烷塔T4顶气相通过主冷箱E4复热至17℃左右进入膨胀/压缩机C1增压端增压至2.0MPag后送至下游外输干气增压机组J1；
外输干气增压机组J1回流约10％的气量作为循环回流干气，在图3中用L3标记出回流干气的管线，回流干气经主冷箱E4冷至-105.2℃后送至脱甲烷塔T4顶部，与脱甲烷塔T4内上升气相进行传质传热，将气相中的C2进一步冷却分离，达到高收率；
为了充分回收脱甲烷塔T4的冷量，脱甲烷塔T4的中部设置两股液相侧抽线LC1和LC2，侧抽线LC1抽出位置在脱甲烷塔T4的34#塔盘、采出温度约-95.5℃，通过主冷箱E4回收冷量后被加热至约-87℃后返回脱甲烷塔T4的下一层塔盘；侧抽线LC2抽出位置在脱甲烷塔T4的40#塔盘、采出温度约-66℃，通过主冷箱E4回收冷量后被加热至约-37℃后返回脱甲烷塔T4的下一层塔盘；两股液相侧抽线LCl和侧抽线LC2的抽出液相量分别为41180kg/h(92.68m3/h)和32840kg/h(67.33m3/h)，抽出量影响塔顶过冷气回流温度，会影响到乙烷收率，本发明的侧线抽出量确保能够获得最大的乙烷收率。
进一步的，如图1、图3所示，为了提升塔底热量，在脱甲烷塔T4的下部设置一股侧抽出线LC3，如图3所示，侧抽出线LC3经主冷箱E4换热预热后回到脱甲烷塔T4下部；三股侧抽出线LC1、LC2和LC3充分利用脱甲烷塔T4物料的阶梯冷量，节能效果明显；
进一步的，在脱甲烷塔T4底再设置脱甲烷塔T4重沸器对原料气预冷同时获得热量，脱甲烷塔T4C2+塔底凝液通过塔底泵P4增压后先经脱甲烷塔T4重沸器复热后再进行凝液分流。
进一步的，主冷箱E4选用8股流板翅式冷箱，最小夹点温度接近3℃，对数平均温差不大于6℃，最大截面温差低于18℃，换热效率高，运行温度应力小，降低了制冷系统负荷，延长了使用寿命；本发明的乙烷回收工艺采用“丙烷预冷、膨胀制冷”组合制冷工艺，合理匹配不同温位制冷需求，节约制冷用量；
主冷箱E4连接有用于补充冷量的丙烷制冷机组；如图5所示，丙烷制冷机组包括丙烷入口分离器J21、丙烷增压机组J22、压缩机出口冷却器J23和丙烷平衡罐J24；混合冷剂制冷机组J3包括冷剂压缩机入口缓冲罐J31、混合冷剂压缩机J32、混合冷剂出口冷却器J33；由于本工艺包膨胀/压缩机C1采用同轴增压膨胀机制冷技术，等熵效率高，运行稳定可靠，膨胀/压缩机C1的膨胀比为约为2.1，制冷负荷不能满足要求，所以主冷箱E4还利用了丙烷冷剂补充冷量，丙烷作为辅助制冷剂，制冷温位为-20℃；来自丙烷制冷机组缓冲的丙烷冷剂经节流降压后进入丙烷平衡罐J24；丙烷平衡罐J24供给主冷箱E4-20℃的液相丙烷，丙烷冷剂在主冷箱E4和丙烷平衡罐J24之间热虹吸循环，丙烷平衡罐J24的气相丙烷返回丙烷制冷机组；本发明的丙烷制冷机组和外输干气增压机组J1可利用旧设备，更加环保。
进一步的，如图4所示，凝液分馏工艺中，乙烷回收单元回收的C2+凝液主要含有乙烷、丙烷、丁烷、戊烷及C5+，为了保证产品乙烷纯度，凝液中的甲烷含量小于0.5％(V/V)。根据凝液中各组分的沸点不同，先后进入脱乙烷塔T5和液化气塔进行分馏得到产品乙烷、LPG及稳定轻烃；具体的，脱甲烷塔T4底来的凝液进入脱乙烷塔T5中部进行分馏，脱乙烷塔T5塔顶气相乙烷在-6.7℃、1.9Mpa条件下进入乙烷冷箱冷凝后进入脱乙烷顶回流罐D51，然后经过脱乙烷顶回流泵P51增压至-14.2℃、2.5Mpa，一部分作为乙烷产品去经乙烷冷箱液化后达到-90℃、2.58Mpa后经节流阀二调节至-90℃、0.02Mpa送至乙烷储罐S1，另一部分回流至脱乙烷塔T5塔顶；乙烷储罐S1中的蒸发气经BOG回收机组J4收回进入脱乙烷顶回流罐D51。脱乙烷塔T5底设置脱乙烷重沸器E51，将塔底温度加热至约70℃；脱乙烷塔T5塔底C3+凝馏送至液化气塔；一台丙烷制冷压缩机组J2提供的丙烷冷剂和乙烷冷箱为脱乙烷塔T5顶提供冷量，制冷温位为-20℃；
来自脱乙烷塔T5的C3+凝液首先与液化气塔底凝液经预热器E52换热后进入液化气塔中部进行分馏，液化气塔塔顶气进入液化气顶冷却器AC61后去液化气顶回流罐D6D6，再经液化气顶回流泵P61增压后一部分作为液化气产品送至罐区，另一部分作为液化气塔塔顶的回流液。液化气塔塔底设置液化气重沸器E61，温度加热至约154℃、压力1.4至1.5MPag；液化气塔塔底稳定轻烃先与C3+凝液经预热器E52换热后再经稳定轻烃冷却器AC63冷却至40℃后送至稳定轻烃球罐S3；液化气经液化气冷却器冷却后送至液化气储罐S2。
进一步的，为了方便产品乙烷的运输和储存，需要对脱乙烷塔T5出的产品乙烷进行过冷至-90℃、0.2Mpag；
乙烷冷箱由混合冷剂和丙烷冷剂提供部分冷量，具体的，混合冷剂制冷机组J3采用螺杆压缩机，混合冷剂压缩机排气压力为1.9Mpa；为了降低能耗，从丙烷制冷机组J2引一股丙烷冷剂至乙烷冷箱，丙烷冷剂先经过丙烷平衡罐二闪蒸分离，液相至乙烷冷箱预冷混合冷剂，丙烷冷剂在出乙烷冷箱后回丙烷平衡罐二。混合冷剂由混合冷剂压缩机增压至1.9Mpa，经混合冷剂出口冷却器J33冷却至约48℃后进入乙烷冷箱，经乙烷冷箱换热至-90℃并经节流阀K5节流降温至-96.70℃、0.32Mpag后返回乙烷冷箱提供冷量，复热后的混合冷剂返回冷剂压缩机入口缓冲罐J31。
本工艺中，主冷箱E4提供约17452.75kW的换热量，丙烷制冷剂提供约1094kW的冷量，膨胀/压缩机C1提供约1453kW的冷量。根据某处气田伴生气气质特点：C2 +组分较贫，约225ml/m3；CO2含量高，约2.2％，利用HYSYS模拟计算软件，针对液体过冷工艺、Ortloff公司部分干气循环工艺以及本发明的工艺进行能耗对比；液体过冷工艺、Ortloff公司部分干气循环工艺以及本发明的工艺进行能耗对比详见下表，表中比较的内容包括从伴生气中提取乙烷的所有步骤，原料气脱碳装置能耗脱碳和乙烷产品脱碳装置能耗是在步骤中为原料气脱碳，原料气增压机组是为进入系统的原料气增压，脱乙烷重沸器E51是脱甲烷塔T4的凝液进入脱乙烷塔T5分离的能耗。
Figure BDA0003299862820000111
由上表可知，本发明的工艺相比于Ortloff公司部分干气循环工艺节约能耗(9.59-8.39)/9.59＝12.5％，节约能耗12.5％，且乙烷收率高，经济效益显著。
尽管参照前述实施例对本发明进行了详细的说明，对于本领域的技术人员来说，其依然可以对前述各实施例所记载的技术方案进行修改，或者对其中部分技术特征进行等同替换，凡在本发明的精神和原则之内，所作的任何修改、等同替换、改进等，均应包含在本发明的保护范围之内。

Claims (5)

1.一种从天然气中提取乙烷的方法，其特征在于，包括以下步骤：
步骤一，原料气增压：将原料气经原料气压缩机增压至4.4MPa、48℃，得到增压原料气；
步骤二，MDEA脱碳：采用活化MDEA溶液吸收增压后的原料气中的CO2，将CO2含量脱除至50ppm以下，得到脱碳后原料气；
步骤三，脱水脱汞：将步骤二中得到的脱碳后原料气，采用三塔等压分子筛脱水脱汞工艺进行脱水和浸硫活性炭脱汞，分子筛脱水脱汞工艺使用的分子筛脱水脱汞装置采用三个干燥塔(T3)和一个小塔，两台干燥塔(T3)在脱水时原料气上进下出，第三台干燥塔(T3)在再生时再生气下进上出，小塔用于再生气的干燥；原料气与再生气通过流量控制实现控制压差＞70kPa，再生气用量约为原料气量的10％；出塔的干气经第三粉尘过滤器(F33)过滤后，得到含水量≤1ppm、汞含量＜0.01μg/m3的净化气；
步骤四，乙烷回收：将步骤三中的净化气经主冷箱(E4)预冷至-64℃至-67℃后进入低温分进行气液分离；低温分离器分离的液相全部送至脱甲烷塔(T4)上部；低温分离器分离的气相分为两股，第一股占比14％-16％的气量经主冷箱(E4)冷至-104℃至-106℃后节流进入脱甲烷塔(T4)上部，第二股经膨胀/压缩机(C1)膨胀至2.0MPag进入脱甲烷塔(T4)上部；脱甲烷塔(T4)塔顶气相通过主冷箱(E4)复热至16℃-18℃进入膨胀/压缩机(C1)增压端增压至2.0MPag后送至外输干气增压机组(J1)；外输干气增压机组(J1)9％-12％的气量作为循环回流干气，回流干气经主冷箱(E4)冷至-104℃至-106℃送至脱甲烷塔(T4)顶部，使得回流的干气能够与脱甲烷塔(T4)内上升气相进行传质传热；脱甲烷塔(T4)塔底凝液经塔底泵(P4)增压后送出；
步骤五，凝液分馏：将脱甲烷塔(T4)底来的凝液送入脱乙烷塔(T5)中部进行分馏，脱乙烷塔(T5)塔顶气相乙烷进入乙烷冷箱冷凝后进入脱乙烷顶回流罐(D51)，然后经过脱乙烷顶回流泵(P51)增压，一部分作为乙烷产品送至乙烷储罐(S1)，另一部分回流至脱乙烷塔(T5)塔顶；
在脱甲烷塔(T4)的中部设置两股液相侧抽出线，两股液相侧抽出线分别经主冷箱(E4)回收冷量后返回脱甲烷塔(T4)的下一层塔盘；脱甲烷塔(T4)的下部设置一股侧抽出线，经主冷箱(E4)换热预热后回到脱甲烷塔(T4)下部；
所述脱甲烷塔(T4)为板式塔，塔盘个数60个，三股液相侧抽出线分别从上向下数第34层塔盘、第40层塔盘和第60层塔盘处抽出；
在步骤五中，乙烷产品送至乙烷储罐(S1)前，通过乙烷冷箱将乙烷产品过冷至-90℃、0.3Mpag；乙烷冷箱采用混合冷剂循环，同时利用丙烷冷剂对混合冷剂预冷，混合冷剂由甲烷、丙烷、乙烯三组分组成，三组分体积比为甲烷16％、丙烷52％、乙烯32％，混合冷剂由混合冷剂制冷机组节流降温至-96.70℃、0.32Mpag后返回乙烷冷箱提供冷量。
2.根据权利要求1所述的从天然气中提取乙烷的方法，其特征在于：步骤二中，MDEA溶液由重量份38.5％的MDEA、56.64％的水、4.86％的哌嗪组成；增压后的原料气送入胺液吸收塔(T21)的底部，自下而上流动的原料气与贫胺液逆流接触，将原料气中的CO2吸收脱除；胺液吸收塔(T21)顶部的净化气经第一冷却器降温、并经净化气分离器分液后进入分子筛脱水脱汞单元进行脱水。
3.根据权利要求1所述的从天然气中提取乙烷的方法，其特征在于：主冷箱(E4)连接有用于补充冷量的丙烷制冷机组，制冷温位-20℃。
4.根据权利要求1所述的从天然气中提取乙烷的方法，其特征在于：步骤五中，脱乙烷塔(T5)底设置脱乙烷重沸器(E51)，将脱乙烷塔(T5)塔底温度加热至70℃，脱乙烷塔(T5)塔底C3+凝馏送至液化气塔。
5.根据权利要求1所述的从天然气中提取乙烷的方法，其特征在于：来自脱乙烷塔(T5)的凝液首先与液化气塔底凝液通过预热器(E52)换热后进入液化气塔中部进行分馏，塔顶气进入液化气顶冷却器后去液化气顶回流罐(D6)，再经液化气顶回流泵增压后一部分作为液化气产品送至罐区，另一部分作为液化气塔顶的回流液；液化气塔塔底设置液化气重沸器，将液化气塔塔底温度加热至154℃、1.4-1.5MPag；液化气塔塔底稳定轻烃先与C3+凝液换热后再经稳定轻烃冷却器(AC63)冷却至40℃后送至储罐。
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