CN1131982C - 用于液化天然气的改进的多组分致冷方法 - Google Patents

用于液化天然气的改进的多组分致冷方法 Download PDF

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CN1131982C
CN1131982C CN988066831A CN98806683A CN1131982C CN 1131982 C CN1131982 C CN 1131982C CN 988066831 A CN988066831 A CN 988066831A CN 98806683 A CN98806683 A CN 98806683A CN 1131982 C CN1131982 C CN 1131982C
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cooling zone
natural gas
logistics
gas
rich
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CN1261952A (zh
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R·R·伯温
E·T·科尔
E·L·基姆伯
E·R·索马斯
L·R·凯利
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ExxonMobil Upstream Research Co
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Exxon Production Research Co
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    • B—PERFORMING OPERATIONS; TRANSPORTING
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    • B23K9/00—Arc welding or cutting
    • B23K9/16—Arc welding or cutting making use of shielding gas
    • B23K9/173—Arc welding or cutting making use of shielding gas and of a consumable electrode
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    • F25J3/02—Processes or apparatus for separating the constituents of gaseous or liquefied gaseous mixtures involving the use of liquefaction or solidification by rectification, i.e. by continuous interchange of heat and material between a vapour stream and a liquid stream
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    • B23K35/24—Selection of soldering or welding materials proper
    • B23K35/30—Selection of soldering or welding materials proper with the principal constituent melting at less than 1550 degrees C
    • B23K35/3053—Fe as the principal constituent
    • B23K35/3066—Fe as the principal constituent with Ni as next major constituent
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    • F25J1/0211—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures requiring the use of refrigeration, e.g. of helium or hydrogen ; Details and kind of the refrigeration system used; Integration with other units or processes; Controlling aspects of the process using a multi-component refrigerant [MCR] fluid in a closed vapor compression cycle
    • F25J1/0219—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures requiring the use of refrigeration, e.g. of helium or hydrogen ; Details and kind of the refrigeration system used; Integration with other units or processes; Controlling aspects of the process using a multi-component refrigerant [MCR] fluid in a closed vapor compression cycle in combination with an internal quasi-closed refrigeration loop, e.g. using a deep flash recycle loop
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    • F25J1/00—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures
    • F25J1/02—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures requiring the use of refrigeration, e.g. of helium or hydrogen ; Details and kind of the refrigeration system used; Integration with other units or processes; Controlling aspects of the process
    • F25J1/0228—Coupling of the liquefaction unit to other units or processes, so-called integrated processes
    • F25J1/0229—Integration with a unit for using hydrocarbons, e.g. consuming hydrocarbons as feed stock
    • F25J1/0231—Integration with a unit for using hydrocarbons, e.g. consuming hydrocarbons as feed stock for the working-up of the hydrocarbon feed, e.g. reinjection of heavier hydrocarbons into the liquefied gas
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    • F25J1/00—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures
    • F25J1/02—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures requiring the use of refrigeration, e.g. of helium or hydrogen ; Details and kind of the refrigeration system used; Integration with other units or processes; Controlling aspects of the process
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    • F25J1/00—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures
    • F25J1/02—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures requiring the use of refrigeration, e.g. of helium or hydrogen ; Details and kind of the refrigeration system used; Integration with other units or processes; Controlling aspects of the process
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    • F25J1/00—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures
    • F25J1/02—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures requiring the use of refrigeration, e.g. of helium or hydrogen ; Details and kind of the refrigeration system used; Integration with other units or processes; Controlling aspects of the process
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    • F25J1/00—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures
    • F25J1/02—Processes or apparatus for liquefying or solidifying gases or gaseous mixtures requiring the use of refrigeration, e.g. of helium or hydrogen ; Details and kind of the refrigeration system used; Integration with other units or processes; Controlling aspects of the process
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    • F17C2205/0302—Fittings, valves, filters, or components in connection with the gas storage device
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    • F17C2260/00—Purposes of gas storage and gas handling
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    • F17C2265/035—Treating the boil-off by recovery with cooling with subcooling the liquid phase
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    • F17C2265/06—Fluid distribution
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    • F17C2265/00—Effects achieved by gas storage or gas handling
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    • F17C2270/0102—Applications for fluid transport or storage on or in the water
    • F17C2270/0105—Ships
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    • F17C2270/00—Applications
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    • F17C2270/0165—Applications for fluid transport or storage on the road
    • F17C2270/0168—Applications for fluid transport or storage on the road by vehicles
    • F17C2270/0171—Trucks
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    • F17C2270/00—Applications
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    • F17C2270/0168—Applications for fluid transport or storage on the road by vehicles
    • F17C2270/0173—Railways
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    • F25—REFRIGERATION OR COOLING; COMBINED HEATING AND REFRIGERATION SYSTEMS; HEAT PUMP SYSTEMS; MANUFACTURE OR STORAGE OF ICE; LIQUEFACTION SOLIDIFICATION OF GASES
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    • F25J2205/00—Processes or apparatus using other separation and/or other processing means
    • F25J2205/02—Processes or apparatus using other separation and/or other processing means using simple phase separation in a vessel or drum
    • F25J2205/04—Processes or apparatus using other separation and/or other processing means using simple phase separation in a vessel or drum in the feed line, i.e. upstream of the fractionation step
    • F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
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    • F25J2205/00—Processes or apparatus using other separation and/or other processing means
    • F25J2205/50—Processes or apparatus using other separation and/or other processing means using absorption, i.e. with selective solvents or lean oil, heavier CnHm and including generally a regeneration step for the solvent or lean oil
    • F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
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    • F25J2210/00—Processes characterised by the type or other details of the feed stream
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    • F—MECHANICAL ENGINEERING; LIGHTING; HEATING; WEAPONS; BLASTING
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Abstract

本发明涉及一种液化富含甲烷的加压气流(10)，以生产温度高于约-112℃(-170)，压力足以使液体产品处于或低于其泡点的富含甲烷的液体产品的方法，其中气流的液化是换热器(33)中进行，换热器由封闭回路多组分致冷系统(45)所冷却。然后，液化气体产品被引入到储存装置(50)在高于约-112℃(-170)的温度下储存。

Description

用于液化天然气的改进的多组分致冷方法
发明领域
本发明涉及一种天然气液化方法，具体地说，涉及一种生产加压液体天然气(PLNG)的方法。
发明背景
由于天然气的干净的燃烧性质和便利性，近年来已被广泛使用。许多天然气源位于遥远的地区，距气体的工业市场有很长的距离。有时，可以利用管线将生产的天然气输送到工业市场。当管线输送不可行时，所生产的天然气通常被加工成液化天然气(以下称“LNG”)以输送到市场。
LNG工厂的一个区别特征是需要大的设备投资。用于液化天然气的设备通常十分昂贵。液化工厂是由几个基本系统组成的，包括处理气体以除去杂质、液化、致冷、动力设备、储存和装货设备。LNG工厂的成本可能随工厂的位置不同变化很大，通常，常规LNG工程可能花费50至100亿美元，包括土地开发成本。工厂的致冷系统可能占到总成本的30％。
在设计LNG工厂时，要考虑的三个主要方面是：(1)选择液化循环，(2)用于容器、管线和其它设备的材料，(3)将天然气进料物流转化成LNG的处理步骤。
由于液化天然气需要很大的制冷量(refrigeration)，所以LNG致冷系统是昂贵的。天然气通常是在约4,830kPa(700psia)至约7,600kPa(1,100psia)的压力和约20℃(68°F)至约40℃(104°F)的温度下进入LNG工厂。天然气主要是甲烷，与用作燃料的重质烃不同，简单加压是不能液化的。甲烷的临界温度是-82.5℃(-116°F)。这意味着不管施加的压力有多高，甲烷只有在这一温度以下才能液化。因为天然气是一种混合物，将在一个温度范围内液化。天然气的临界温度在约-85℃(-121°F)至-62℃(-80°F)之间。通常，在常压下，天然气组合物在约-165℃(-265°F)至-155℃(-247°F)液化。因为致冷设备占据了LNG设备成本中很重要的一部分，所以，人们致力于降低致冷成本。
尽管已有许多致冷循环用于液化天然气，但目前用于LNG工厂的三种典型循环是：(1)“串级(cascade)循环”，在依次设置的换热器中使用多个单组分致冷剂以将气体温度降低到液化温度，(2)“膨胀机循环”，将气体从高压膨胀到低压，相应地降低温度，(3)“多组分致冷循环”，在特殊设计的换热器中使用多组分致冷剂。大多数天然气液化循环使用这三种基本循环的改进或结合。
混合致冷剂系统包括循环多组分致冷物流，通常是在用丙烷预冷却至约-35℃(-31°F)以后。典型的多组分致冷剂系统可以含有甲烷、乙烷、丙烷以及任选的其它轻质烃。如果不使用丙烷预冷却，则可以在多组分致冷剂中包括如丁烷和戊烷的重质烃。混合致冷剂循环的特性应当使过程中的换热器常规地处理两相致冷剂的流动，这就要求使用大的特定换热器。混合致冷剂表现出在一定温度范围内冷凝的性质，这使得难以设计出在热动力学上比纯组分致冷剂系统更有效的换热器系统。多组分致冷方法的例子公开在US5502972；5497626；3763638和4586942中。
用于常规LNG工厂的材料也计算在工厂的成本内。用于LNG工厂的容器、管线、其它设备通常至少一部分是由铝、不锈钢、或高镍含量的钢构成，以在低温下提供必要的强度和断裂韧性。
在常规LNG工厂中，水、二氧化碳、如硫化氢和其它酸性气体的含硫化合物、正戊烷和包括苯的重质烃必须从天然气加工过程中除去，以达到百万分之几(ppm)的水平。这些化合物中的某些物质会冻结，从而在处理设备中引起堵塞问题。其它化合物，如含硫化合物，通常被除去以满足出售的规定。在常规的LNG工厂中，需要气体处理设备以除去二氧化碳和酸性气体。气体处理设备通常是使用化学和/或物理溶剂再生方法，需要投资大量资金。此外，操作费用也高。需要如分子筛的干燥床脱水器以除去水蒸汽。通常使用洗涤塔和分馏设备以除去可能产生堵塞问题的烃。由于汞会引起由铝构造的设备的故障，因此，在常规LNG工厂中也要除去。此外，在处理后，可能存在于天然气中的大部分氮气也要除去，因为在传统LNG输送过程中，氮气不会保留在液相中，从输送的观点看，在LNG容器中存在氮蒸汽是所不希望的。
在工业上需要一种改进的天然气液化方法，以使致冷设备和所需的能耗最小。
发明概述
本发明涉及一种富含甲烷的气流的液化方法。进料气流的压力高于约3,100kPa(450psia)。如果压力太低，可以先压缩气体。通过利用多组分致冷系统液化气体，以产生温度高于约-112℃(-170°F)和压力足以使液体产品处于或低其泡点的液体产品，产品在这里称之为加压液体天然气(“PLNG”)。在用多组分致冷液化之前，优选使用通过膨胀装置的、但未液化的循环蒸汽冷却气体。在高于约-112℃(-170°F)的温度下将PLNG引入储存装置以储存。
在本发明的另一方案中，如果进料气体含有重于甲烷的组分，可以在通过多组分致冷液化之前，通过分馏方法除去重质烃。
在本发明的再一方案中，由液化天然气蒸发所产生的沸腾气体可以加入到进料气体中，通过多组分液化，以产生PLNG。
本发明方法可以用于在供应储存或输送天然气的源地最初液化天然气，也可以用于储存和装货时产生的天然气蒸汽的再液化。因此，本发明的一个目的是液化或再液化天然气的改进液化系统。本发明的另一目的是提供一种改进的液化系统，所需的压缩能量比现有系统少得多。本发明的再一目的是提供一种操作经济而有效的改进液化系统。与本发明的生产PLNG的相对中等温度致冷相比，在非常低的温度下致冷的常规LNG方法是十分昂贵的。
附图简述
参照下面的详细说明和附图，可以更好地理解本发明的优点，其中附图是说明本发明实施方案的流程示意图。
图1是本发明生产PLNG的第一方案的流程示意图，示出了生产PLNG的封闭回路多组分致冷系统。
图2是本发明第二方案的示意流程图，其中，在液化成PLNG之前，天然气被分馏。
图3是本发明第三方案的流程示意图，其中，在液化成PLNG之前，用封闭单组分致冷系统预冷却天然气。
图4是本发明第四方案的示意流程图，其中，在分馏之前，封闭回路多组分系统预冷却天然气流，致冷系统还将天然气流液化成PLNG。
图5是本发明第五方案的示意流程图，其中，天然气被分馏，然后，在由第二封闭回路致冷系统冷却的换热器中液化，第二封闭回路致冷系统使用多组分液体和多组分蒸汽两者作为致冷剂。沸腾蒸汽仅使用多组分致冷系统的蒸汽来液化。
图6是本发明第六方案的示意流程图，其中，在被多组分致冷系统液化成PLNG之前，沸腾蒸汽与天然气进料混合。
图7是本发明第七方案的示意流程图，其中，天然气被分馏，然后，在由第二封闭回路致冷系统冷却的换热器中液化，第二封闭回路致冷系统使用多组分液体和多组分蒸汽两者作为致冷剂。
图8是用于图2、5、6和7所示方案的膨胀机过程
图9是用于图1、2、3、4和6所示方案的优选多组分致冷系统的示意流程图。
图10是用于图5和7所示方案的优选多组分致冷系统的示意流程图。
附图中的流程图代表了实施本发明方法的不同方案。附图不排除本发明范围内的其它实施方案，它们可以是这些特定方案的正常或预期改进的结果。为了简化和清楚的目的，从附图中省略了所需的各种次级系统，如泵、阀、物流混合器、控制系统、传感器等。优选方案的描述
本发明涉及一种使用多组分致冷系统液化天然气以产生富含甲烷的液体产品，其温度高于约-112℃(-170°F)，压力足以使液体产品处于或低于泡点。在本发明的描述中，富含甲烷的产品有时被称之为加压液体天然气(PLNG)。术语“泡点”是指液体开始转变成气体的温度和压力。例如，如果一定体积的PLNG保持在恒定的压力下，当其温度升高时，开始在PLNG中形成气泡的温度就是其泡点。类似地，如果一定体积的PLNG保持在恒定的温度下，当压力下降时，开始形成气泡的压力定义为泡点。在泡点下，混合物是饱和液体。
使用本发明的多组分致冷系统液化天然气所需要的能量比现用方法少，用于本发明的设备可以用较便宜的材料制成。与此相反，现有方法在常压下生产温度低至-160℃(-256°F)的LNG，要求处理设备由昂贵的材料制成以保证安全。
相对于常规LNG工厂所需的总能量，在本发明的实施中，液化天然气所需能量大大减少。本发明方法所需致冷能量的下降导致了资金成本的大幅度下降，成比例地降低了操作费用，提高了效率和可靠性，因此，大大增强了生产液化天然气的经济性能。
在本发明的操作压力和温度下，在液化过程的最冷操作区域的管线和设备中，可以使用约3.5％重量的镍钢，而在常规LNG过程中，需要使用更贵的9％重量的镍或铝钢。这样，与现有LNG方法相比，提供了另一明显的成本下降。
在低温天然气处理中首先要考虑的是污染问题。适合于本发明方法的粗天然气原料可以是从粗油井(结合气)或气井(非结合气)获得的天然气。天然气的组成可能在大范围内变动。在这里，天然气物流含有甲烷(C1)作为主要组分。天然气通常可以含有乙烷(C2)、重质烃(C3+)，以及少量的污染物，如水、二氧化碳、硫化氢、氮气、丁烷、六个或更多碳原子的烃、泥土、硫化铁、石蜡、原油。这些污染物的溶解度随温度、压力和组成而变化。在低温下，CO2、水和其它污染物可能形成固体，堵塞低温换热器中的流动通道。如果条件在它们的纯组分范围内，固相温度-压力相边界提前出现，通过除去这些污染可以避免这些潜在的困难。在本发明的以下描述中，假定天然气已经过适当的处理以除去硫化物和二氧化碳，并已使用常规的已知方法干燥以除去水分，产生了“完美、干燥”的天然气物流。如果天然气物流含有会在液化过程中冻结的重质烃，或如果重质烃在PLNG中不所不希望的，则重质烃可以在生产PLNG之前通过分馏方法除去，这将在下面详细描述。
本发明的一个优点是较高的操作温度，保证了天然气具有高于常规LNG方法的可冻结组分浓度。例如，在常规LNG工厂中，在-160℃(-256°F)的温度下生产LNG，CO2必须低于约50ppm以避免堵塞问题。与此相反，通过将过程的温度保持在约-112℃(-170°F)以上，在-112℃(-170°F)时，天然气可以含有高至约1.4摩尔％的CO2，在-95℃(-139°F)时，可以含有约4.2摩尔％的CO2，在本发明的液化过程中不会发生冻结问题。
此外，在本发明方法中，天然气中含有中等量的氮气不必除去，因为在本发明的操作压力和温度下，氮气与液化的烃保持在液相状态。当天然气的组成允许时，减少或省略气体处理和排出氮气所需的设备，提供了很大的技术和经济上的优点。参看附图，可以更好地理解本发明的这些和其它优点。
参看图1，加压天然气进料物流10优选在高于约1,724kPa(250psia)，进一步优选高于约4,827kPa(700psia)的压力和优选低于约40℃(104°F)的温度下进入液化过程；然而，如果必要的话，可以使用不同的压力和温度，考虑到本发明的教导，本领域内的技术人员可以适当改进系统。如果气体物流10低于约1,724kPa(450psia)，可以使用适当的压缩装置(未示出)进行压缩，压缩装置可以包括一个或多个压缩机。
天然气进料物流10被输送到进料冷却器26中，将天然气流冷却到低于约30℃(86°F)的温度，该冷却器可以是任何常规冷却系统。冷却优选通过换热器用空气或水进行。从进料冷却器26排出的冷却物流11被输送到常规多组分换热器33的第一冷却区33a，这种换热器可以从市场上购买，是本领域技术人员所熟知的。本发明不限于任何种类的换热器，但是由于经济原因，优选板-翅式、盘管和低温度箱换热器。优选地，被输送到换热器中的含有液相和气相的所有物流在它们进入的通道的横截面积上均匀分布液相和气相。为了实现这一点，优选对蒸汽和液体物流分别提供分配设备。可以在多相物流中添加分离器将物流分离成液相和汽相。例如，可以在物流18和24进入冷却区33a和33b之前，将分离器分别上加入到图1所示物流18和24上(图1未示出这种分离器)。
换热器33具有一个或多个冷却区，优选至少两个冷却区。图1所示换热器33具有两个冷却区33a和33b。由于与来自多组分致冷系统45的致冷剂换热，天然气物流11在冷却区33a中被液化，多组分致冷系统45在本说明书中称之为MCR系统45。MCR45的优选方案如图9所示，将在下面详细讨论。MCR系统中的致冷剂由烃的混合物制成，例如，包括甲烷、乙烷、丙烷、丁烷和戊烷。优选致冷剂的摩尔百分比组成为：甲烷(25.8％)，乙烷(50.6％)，丙烷(1.1％)，异丁烷(8.6％)，正丁烷(3.7％)，异戊烷(9.0％)，和正戊烷(1.2％)。可以调节MCR组分的浓度，使之与要冷却的进料气体的冷却和冷凝特性以及液化过程所需的低温相匹配。作为适合于封闭回路MCR致冷系统的温度和压力的例子，管线27中的345kPa(50psia)和10℃(50°F)的致冷剂可以直接引入到MCR系统45中以进行常规压缩和冷却，产生压力为1,207kPa(175psia)、温度为13.3℃(56°F)的多组分物流18。物流18在冷却区33a中冷却，在冷却区33b中进一步冷却，从冷却区33b中排出的冷物流23的温度为-99℃(-146°F)。然后，物流23经常规焦耳-汤姆逊阀46膨胀，产生414kPa(60psia)和-108℃(-162°F)的物流24。然后，物流24在冷却区33b中升温，产生10℃(50°F)和345kPa(50psia)的物流27。然后，多组分致冷剂在封闭回路致冷系统中循环。在图1所示液化过程中，MCR系统45是用于生产PLNG的唯一封闭回路致冷系统。
液化的天然气流19是温度高于约-112℃(-170°F)压力足以使液体产品处于或低于其泡点的PLNG。如果物流19的压力高于保持物流10为液相的压力，则物流19可以任选地通过一个或多个膨胀装置，如水轮机34，以产生温度仍高于约-112℃(-170°F)压力足以使液体产品处于或低于其泡点的PLNG。然后，PLNG通过管线20和29输送到适合的储存或运输装置，如管线、储罐，或如PLNG船、罐车或有轨车的输送器。
在储存、运输或处理液化天然气的过程中，可能存在大量“沸腾蒸汽”，这些蒸汽是由液化天然气蒸发产生的。本发明特别适合于液化由PLNG产生的沸腾蒸汽。本发明方法可以任选地再液化沸腾蒸汽。参看图1，沸腾蒸汽通过管线22引入本发明的方法。物流22中的一部分任选地被引导到冷却区33a，以使排出的沸腾气体升温，并向冷却区33a提供附加的制冷量，升温后的排放气体用作燃料。物流22中剩余的一部分被输送到冷却区33b，沸腾气体在其中被液化。从冷却区33b排出的液化天然气(物流28)由泵36加压至水轮机34排出的PLNG的压力，然后，与物流20合并，被输送到合适的储存50中。
从水轮机34和泵36排出的物流优选通过一个或多个相分离器(在图中未示出这种分离器)，从液化天然气中分离出在过程中未被液化的任何气体。这种分离器的操作对本领域技术人员来说是熟知的。然后，液化气体被输送到PLNG储存装置50，来自相分离器的气相可以用作燃料或循环到过程以液化。
图2示出了本发明的另一种方案，在说明书的这一附图和其它附图中，用相同数字表示的部件具有相同的功能。然而本领域技术人员可以认识到，处理设备从一方案到另一方案可以在尺寸和容量上发生变化以处理不同的流体流量、温度和组成。参看图2，天然气进料物流经管线10进入系统，通过常规进料冷却器26。来自进料冷却器26的天然气被输送到膨胀机过程30中，膨胀机过程将天然气流冷却到足以使天然气中至少大部分重质烃组分冷凝的温度，称之为天然气液体(NGL)。NGL包括乙烷、丙烷、丁烷、戊烷、异戊烷等。在4,137kPa(600psia)至7,585kPa(1,100psia)的压力下，实现冷凝所需的温度为约0℃(32°F)至约-60℃(-76°F)。膨胀机过程30的优选方案如图8所示，这将在下面详细描述。
来自膨胀机过程30的塔底物流12被输送到常规分馏设备35中，其操作对本领域技术人员来说是熟知的。分馏设备35可以包括一个或多个分馏塔(在图2中未示出)，将液体塔底物流12分离成主要含乙烷、丙烷、丁烷、戊烷和己烷的物流。分馏设备优选包括多个分馏塔(未示出)，如产生乙烷的脱乙烷塔，产生丙烷的脱丙烷塔，产生丁烷的脱丁烷塔，所有这些组分都可以用于多组分致冷系统45或任何其它合适致冷系统的补充致冷剂。补充致冷剂如图2中物流15所示。如果进料物流10含有高浓度的CO2，可能需要处理一股或多股致冷剂补充物流15，以除去CO2，从而避免在致冷设备中发生潜在的堵塞问题。如果致冷剂物流中CO2的浓度超过3摩尔％，分馏设备35优选包括除去CO2的过程。从分馏设备35中排出液体作为冷凝产物，如图2中物流14所示。从分馏设备35的分馏塔排出的塔顶物流富含乙烷和轻质轻，如图2的物流13所示。
从脱甲烷塔30排出的富含甲烷的物流16与富含乙烷的物流13合并，作为物流17被输送到混合致冷剂冷却区33a以液化天然气。冷却区33a的制冷量由常规多组分致冷系统45提供，已在上面参照图1的MCR系统进行了详细的描述。尽管MCR致冷剂在封闭的回路中循环，但是，如果致冷剂由于从系统中泄漏而受到损失，可以从分馏设备35得到补充(物流15)。在图2所示的的液化过程中，多组分致冷封闭回路45是用于液化天然气进料气流的唯一封闭回路致冷系统。
从混合致冷剂冷却区33a排出的液化天然气流19被输送通过水轮机34降低压力，以在高于约-112℃(-170°F)的温度和足以使液体产品处于或低于其泡点的压力产生PLNG。这一方案的主要优点是可以在膨胀设备中除去重质烃，并可以从分馏设备35中补充致冷剂。
图3示出了本发明的另一方案，其中，在液化成PLNG之前，用单组分致冷系统预冷却天然气流10。图3所示过程类似于图2所示过程，例外的是使用封闭循环致冷系统40提供进料冷却器26所需的至少一部分制冷量，还提供换热器60所需的制冷量。从进料冷却器26排出的物流11被直接输送到脱甲烷塔80，而不需用于图2所示过程的膨胀机过程。致冷系统40可以是使用丙烷、丙烯、乙烷、二氧化碳或其它任何合适液体作为致冷剂的常规系统。
在图3中，来自MCR系统45的管线18a中的液体致冷剂可任选地在换热器70中由物流27中的致冷剂所冷却，物流27是从换热器33返回MCR系统45。物流18a可以在换热器60中由来自致冷系统40中的致冷剂所冷却，致冷系统40具有一股致冷剂物流51在致冷系统40与换热器60之间循环。在这一方案中，大部分冷却致冷剂移到了常规的、纯组分的、封闭回路致冷系统40中，如丙烷致冷系统中。尽管需要附加的换热器，但是换热器33的尺寸和成本降低了。
图4示出了本发明方法的再一方案，其中，封闭回路多组分致冷系统33在分馏之前预冷却了天然气进料物流，致冷系统还液化天然气流产生PLNG。天然气进料物流经管线10进入系统，被输送通过进料冷却器26，它可冷却并部分液化天然气。然后，天然气经管线11被输送到多组分换热器33的第一冷却区33a。在这一方案中，换热器具有三个冷却区(33a，33b，33c)。第二冷却区33b位于第一冷却区33a和第三冷却区33c之间，在低于第一冷却区但高于第三冷却区的温度下操作。
部分液化的天然气排出第一冷却区33a，经管线11a输送到脱甲烷塔80。脱甲烷塔80分馏天然气以产生富含甲烷的塔顶物流16和塔底物流12。底塔物流12被输送到分馏设备35中，该分馏设备与前面图2中所示设备类似。来自脱甲烷塔30的富含甲烷的物流16与来自分馏设备35的塔顶产品物流13合并，并作为物流17被输送到换热器33的第二冷却区33b中。从第二冷却区33b排出的物流19被输送通过如水轮机34的一个或多个膨胀装置。水轮机34产生冷的、膨胀的物流20(PLNG)，在高于约-112℃(-170°F)的温度和足以使液体产品处开或低于其泡点的压力下被输送到储存装置50。
液化天然气在运输或装载过程中在储存器中蒸发产生的沸腾气体可以任选地通过管线22引入到第三冷却区33c，在该冷却区沸腾气体被液化。任选地，一部分沸腾气体可以输送通过第二冷却区33b以加热沸腾气体，然后，作为燃料(物流38)。从冷却区33c排出的液化天然气通过泵36加压到物流20中PLNG的压力，并输送到储存装置50。
图4所示允许除去重质烃和补充致冷剂，而不需图2方案所需的明显的压力降，以及图3所示方案的附加致冷系统。
图5示出了本发明的又一方案，其中，进料天然气被进料冷却器26所冷却，天然气在由封闭回路致冷系统45冷却的换热器33中被液化，致冷系统45使用多组分液体和多组分蒸汽两者作为致冷剂。这样，允许使用唯一的多组分蒸汽来液化储罐中的沸腾蒸汽。这一方案与图2所示方案类似，不同的是多组分换热器33的操作。使用蒸汽和液体致冷剂的MCR系统45的优选方案如图10所示，这将在下面详细讨论。
参看图5，天然气进料物流通过管线10进入系统，被输送通过包括一个或多个部分液化天然气的换热器的进料冷却器26。在该方案中，优选用空气或水通过换热器进行冷却。进料冷却器26任选由常规封闭回路致冷系统40冷却，其中的冷却致冷剂是丙烷、丙烯、乙烷、二氧化碳或其它任何合适的致冷剂。
作为适合于图5所示封闭回路MCR系统45的温度和压力的例子，管线27中的345kPa(50psia)和10℃(50°F)下的多组分致冷剂，被引导到常规MCR系统45中压缩和冷却，产生多组分液体物流18和多组分蒸汽物流21，每物流的压力为1207kPa(175psia)，温度为13.3℃(56°F)。蒸汽物流21在冷却区33a中进一步冷却，在冷却区33b中更进一步冷却，产生冷物流23，排出冷却区33b时的温度为-99℃(-146°F)。然后，物流23通过常规焦耳-汤姆逊阀46膨胀，产生414kPa(60psia)和-108℃(-162°F)的物流24。然后，物流24在冷却区33b中升温，在冷却区33a中进一步升温，在10℃(50°F)和345kPa(50psia)下产生物流27。物流18在冷却区33a中冷却，然后，通过常规焦耳-汤姆逊阀47膨胀。从膨胀阀47排出的膨胀流体与物流25合并，并继续循环。这一方案的优点是使用唯一的MCR蒸汽再液化沸腾蒸汽。
图6示出了本发明的再一方案，它类似于图2所示方案，不同的是多组分换热器33只有一个冷却区(33a)，而沸腾蒸汽与天然气物流16和13混合，而不是用换热器33的单独的冷却区液化。沸腾蒸汽22首先通过冷却区33a，向通过换热器33a的温度较高的物流17和18提供制冷量。从冷却区33a排出后，可以优选地排放物流22中的一部分作为燃料，向PLNG工厂提供能量。物流22中剩余的一部分通过压缩机39，将沸腾气体加压到接近物流17中气体的压力。然后，从压缩机39中排出的沸腾气体(物流32)与物流17合并。这一方案不需要混合低温液体，与图2所示方案相比，可以简化操作。
图7描述了本发明的又一方案，其中进料气体由进料冷却器26冷却，天然气在由封闭回路致冷系统45冷却的多组分换热器33中液化，该致冷系统使用多组分液体(物流18)和多组分蒸汽(物流21)作为致冷剂。图7所示处理过程与图5所示方法的操作类似，不同的是沸腾蒸汽22中的至少一部分被压缩机39压缩到接近气流16的压力，压缩的沸腾物流32与天然气流16合并。然后，含有来自膨胀机过程30的蒸汽、来自分馏设备35的蒸汽和来自物流32的沸腾蒸汽的物流17通过换热器33的冷却区33a和33b，将气流液化产生PLNG(物流19)。图7，优选排出物流22中的一部分，使之通过换热器区33b和33a，作为燃料排出换热器33(物流38)。
用于本发明图2、5、7和7所示方案的优选膨胀机过程如图8所示。参看图8，物流11分成两股100和101。物流100在换热器102中被管线104中的冷残余气体冷却。物流101被侧再沸换热器105冷却，来自脱甲烷130的脱甲烷液体物流通过该换热器。冷却的物流110和101再合并，合并的物流103被输送到常规相分离器106中。分离器106将物流103分成液体物流107和蒸汽物流108。蒸汽物流108通过如涡轮膨胀机109膨胀以降低其压力。在其进入脱甲烷塔80上部之前，这一膨胀进一步冷却了气流。冷凝液体物流107通过焦耳-姆逊阀110膨胀，在其通过脱甲烷塔80之前，进一步冷却物流107。
来自脱甲烷塔80的残余气体被输送到换热器102，并通过压缩机111，压缩机111的至少一部分能量由膨胀机109提供。从膨胀机过程30排出的、压缩的富含甲烷的物流106按本发明方法进一步处理。脱甲烷塔产生了一股塔底液体物流12，主要是天然气液体(NGL)，含有乙烷、丙烷、丁烷、戊烷和重质烃。可用于本发明的另一膨胀机过程30公开在US4698081和《气体调节和加工》，先进技术及应用，第三卷，John M.Campbell和Co.，Tulsa，Oklahoma(1982)中。
图9描述了用于本发明图1、2、3、4和6所示方案的优选MCR系统45的示意流程图。参看图9，物流27进入常规压缩机150以压缩致冷剂。在其进入常规相分离器153之前，来自压缩机150的压缩物流151通过常规冷却器153进行冷却，如空气或水冷却器。来自相分离器153的的蒸汽154被输送到压缩机155。来自压缩机155的压缩致冷剂蒸汽(物流156)由常规冷却器157冷却以产生冷却的致冷剂物流18。来自相分离器152的液体物流158被泵159加压至接近压缩机155出口压力。在被冷却器107冷却之前，来自泵159加压液体(物流160)与物流156合并。
图10是用于图5和7所示方案的优选MCR系统45的示意流程图。图10所示MCR系统类似于图9所示MCR系统45，不同的是液体致冷剂物流160与蒸汽物流156合并，并在冷却器157中冷却后，来自冷却器157的冷物流被输送到常规相分离器161中。从分离器161排出的蒸汽成为蒸汽物流21，从分离器161中排出的液体成为液体物流18。
实施例
模拟质量和能量平衡以说明附图所示的实施方案，其结果列于表1-7中。表中的数据用于更好地理解附图1-7中所示方案，但本发明不受其不必要的限制。表中的温度和流量不是对本发明的限定，从这里所公开的技术来看，温度和流量可以有许多的变化。表与附图的对应关系为：表1对应于图1；表2对应于图2；表3对应于图3；表4对应于图4；表5对应于图5；表6对应于图6；表7对应于图7。
数据是使用市售过程模拟程序HYSYSTM获得的，然而，也可以使用其它市售过程模拟程序来获得数据，包括HYSIMTM、PROIITM和ASPENPLUSTM，对本领域的普通技术人员来说都是很熟悉的。
表3中的数据是假定图3所示方案是用丙烷致冷系统40冷却进料物流10。
使用图3所示基本流程，并使用相同的进料组成和温度，(在接近常压和-160℃(-256°F)的温度下)生产LNG所需的能量是按图3所示方案生产PLNG所需通量的两倍多：生产LNG为185,680kW(249,000hp)，生产PLNG为89,040kW(119,400hp)。这一比较是通过HYSIMTM过程模拟器进行的。
本领域的技术人员，特别是获得了本专利教导的技术人员，将体会到上述方法的许多改进和变动。例如，根据系统的总体设计和气体进料的组成，按照本发明可以改变温度和压力。气体进料的冷却排列方式(train)可以根据总体设计进行增补和重新构造，以满足最优而有效的换热要求。正如上面所讨论的，所公开的特定方案和实施例不应用于限制本发明的范围，这一范围由下面的权利要求及其等同物所确定。
表1
    相            压力             温度           流量                   组成，摩尔％
 物流 汽(V)/液(L)     kPa     psia     ℃     °F  kg摩尔/hr   lb摩尔/hr    C1    C2     C3+   CO2    N2
  10111819202223242527282938     VVV/LLLVLV/LV/LVLLV     5,5715,5022,0685,2952,5862,5861,7394143793452,4482,5862,517     808798300768375375260605550355375365     21.113.313.3-93.9-95.6-94.4-103.3-106.1-99.411.1-103.3-95.611.1     705656-137-140-138-154-159-14752-154-14052     7172717296137172717279796139613961396134287620384     15,81315,81321,19315,81315,8131,75721,19321,19321,19321,19394416,801846   84.0684.062084.0684.0698202020209884.8598   6.876.87466.876.870.68464646460.686.520.68     8.68.6348.68.60.09343434340.098.110.09   0.180.1800.180.180.0500000.050.180.05   0.290.2900.290.291.1800001.180.341.18
               表1(续)
               能耗
能耗hp 能耗kW
压缩机MCR致冷系统45的压缩机第1级第2级膨胀机膨胀机34泵泵36MCR致冷系统45的泵净能耗总设备能耗 13,8004,700-2702110 10,2913,505-201182
18,30018,900 13,64714,094
表2
    相            压力               温度             流量                         组成，摩尔％
    物流 汽(V)/液(L)     kPa     psia     ℃     °F  kg摩尔/hr  lb摩尔/hr    C1   C2     C3+   CO2    N2
    10111213141516171819202223242527282938     V/LV/LLV/LLV/LVVV/LLLVLV/LV/LVLLV     5,5855,514,7575,0191383,3785,0195,0192,0684,8132,8612,8271,7934143793452,6892,8612,758     81080069072820490728728300698415410260605550390415400     21.14.4206.762.226.713.35.04.413.3-93.3-95.6-90.0-99.4-108.3-104.410.0-99.4-95.610.0     70404041448056414056-136-140-130-147-163-15650-147-14050     36,70736,7076821155085536,01036,15952,04836,15936,1592,98852,04852,04852,04852,0481,58437,7031,410     80,92980,9291,5042531,12012179,39279,721114,75079,72179,7216,589114,750114,750114,750114,7503,49283,1253,108   92.692.63.9711.21025.8194.2794.0225.8194.0294.0299.1125.8125.8125.8125.8199.1194.2399.11  3.93.99.5432.57050.633.793.8850.633.883.880.4650.6350.6350.6350.630.463.740.46     2.482.4885.4453.2510023.560.921.0823.561.081.080.0123.5623.5623.5623.560.011.030.01   0.980.981.052.97000.9800.980.980.2800000.280.960.28   0.040.0400000.0400.040.040.1400000.140.040.14
                 表2(续)
                 能耗
能耗hp 能耗kW
压缩机膨胀机过程30的压缩机MCR致冷系统45的压缩机第1级第2级分馏设备35的再压缩机膨胀机膨胀机过程30的膨胀机膨胀机34泵泵36MCR致冷系统45的泵分馏设备35的产品泵净能耗总设备能耗 2,30075,00028,00010-2,300-1,0501048020 1,71555,92820,8807-1,715-783735815
102,500109,200 76,43581,432
表3
    相         压力             温度            流量                        组成，摩尔％
   物流 汽(V)/液(L)    kPa     psia     ℃     °F  kg摩尔/hr  lb摩尔/hr    C1    C2     C3+   CO2    N2
    10111213141516171818a1920222324252727a282938     V/LV/LLV/LLV/LVV/LV/LVLLVLV/LV/LVVLLV   5,5165,3785,3785,2951383,3785,2955,2952,5862,7235,0882,8612,8272,3104143793452762,6892,8612,758     8007807807682049076876837539573841541033560555040390415400     4.4-34.4187.861.726.713.3-34.4-33.9-34.413.3-92.8-95.6-90.0-99.4-107.8-103.3-35.67.8-99.495.6-35.6     40-303701438056-30-29-3056-135-140-130-147-162-154-3246-147-140-32     36,70736,7078171695489035,91036,06026,99526,99536,06036,060298826,99526,99526,99526,99526,995142937,5041,559     80,92980,9291,8013731,20819879,17279,50259,51759,51779,50279,5026,58959,51759,51759,51759,51759,5173,15282,6863,437   92.692.65.4312.33026.0594.5894.2252594.294.299.11252525252599.1194.3899.11   3.93.913.0433.85054.693.693.8375753.833.830.4675757575750.463.70.46     2.482.4880.0550.4710019.260.720.95000.950.950.01000000.010.930.01   0.980.981.483.35000.970.98000.980.980.28000000.280.950.28   0.040.0400000.040.04000.040.040.14000000.140.040.14
               表3(续)
                能耗
能耗hp 能耗kW
压缩机致冷系统40的压缩机第1级第2级MCR致冷系统45的压缩机第1级第2级分馏设备35的再压缩机膨胀机膨胀机34泵泵36分馏设备35的产品泵净能耗总设备能耗 14,60029,70052,70021,10020-1,2001025 10,88722,14839,29915,73515-895719
117,000119,400 87,24889,038
表4
    相       压力            温度              流量                     组成，摩尔％
    物流 汽(V)/液(L)     kPa   psia     ℃     °F  kg摩尔/hr  lb摩尔/hr    C1   C2     C3+   CO2    N2
    101111a12131415161819202223242527282938     V/LV/LV/LLV/LLV/LVV/LLLVLV/LV/LVLLV     5,5855,5165,3785,3785,2951383,3785,2952,7585,0882,8612,8272,2754143793102,6892,8612,758     81080078078076820490768400738415410330605545390415400     21.14.4-34.4187.861.726.713.3-34.413.3-92.8-95.6-90.0-99.4-108.3-104.411.7-99.4-95.6-41.1     7040-303701438056-3056-135-140-130-147-163-15653-147-140-42     36,70736,707,36,7078171695489035,91043,33136,06036,0602,98843,33143,33143,33143,3311,58437,6541,405     80,92980,92980,9291,8013731,20819879,17295,53479,50279,5026,58995,53495,53495,53495,5343,49283,0163,097   92.692.692.65.4312.33026.0594.5826.2594.294.299.1126.2526.2526.2526.2599.1194.499.11  3.93.93.913.0433.85054.693.6950.53.833.830.4650.550.550.550.50.463.690.46     2.482.482.4880.0550.4710019.260.7223.250.950.950.0123.2523.2523.2523.250.010.920.01   0.980.980.981.483.35000.9700.980.980.2800000.280.950.28   0.040.040.0400000.0400.040.040.1400000.140.040.14
             表4(续)
              能耗
能耗hp 能耗kW
压缩机MCR致冷系统45的压缩机第1级第2级分馏设备35的再压缩机膨胀机膨胀机34泵泵36MCR致冷系统45的泵分馏设备35的产品泵净能耗设备总能耗 70,50031,90020-12,001067025 52,57323,78815-895750019
101,900104,300 75,98877,778
表5
    相           压力             温度               流量                       组成，摩尔％
    物流 汽(V)/液(L)     kPa     psia     ℃     °F  kg摩尔/hr  1b摩尔/hr    C1  C2   C3+   CO2    N2
    1011121314151617181920212223242527282938     V/LV/LLV/LLV/LVVLLLVVLV/LV/LVLLV     5,5855,5164,7575,0191383,3785,0195,0192,0684,8132,8612,0682,8271,7934143793452,6892,8612,758     81080069072820490728728300698415300410260605550390415400     21.14.4206.762.226.713.35.04.417.8-93.3-95.617.8-90.0-99.4-113.9-110.69.4-99.4-95.69.4     70404041448056414064-136-14064-130-147-173-16749-147-14049     36,70736,7076821155085536,01036,15918,82736,15936,15940,3432,98840,34340,34340,34359,1701,43437,5541,559     80,92980,9291,5042531,12012179,39279,72141,50879,72179,72188,9456,58988,94588,94588,945130,4533,16382,7963,437   92.692.63.9711.21025.8194.2794.024.8194.0294.0233.6999.1133.6933.6933.6924.599.1194.2199.11  3.93.99.5432.57050.633.793.8837.973.883.8857.670.4657.6757.6757.6751.40.463.750.46   2.482.4885.4453.2510023.560.921.0857.221.081.088.640.018.648.648.6424.10.011.040.01   0.980.981.052.97000.980.9800.980.9800.2800000.280.960.28   0.040.0400000.040.0400.040.0400.1400000.140.14
                  表5(续)
                   能耗
能耗hp 能耗kW
压缩机膨胀过程30的压缩机MCR致冷系统45的压缩机第1级第2级分馏设备35的再压缩机膨胀机膨胀过程30的膨胀机膨胀机34泵泵36MCR致冷系统45的泵分馏设备35的产品泵净能耗设备总能耗 2,30084,90031,80010-2,300-1,0501050020 171563,31123,7147-1,715-783737315
116,200122,900 86,65291,648
表6
    相          压力              温度             流量                        组成，摩尔％
   物流 汽(V)/液(L)   kPa     psia     ℃     °F  kg摩尔/hr  lb摩尔/hr    C1    C2     C3+   CO2    N2
    1011121314151617181920222324273238     V/LV/LLV/LLV/LVVV/LLLVLV/LVVV   5,5855,5164,7575,0191383,3785,0195,0192,0684,8132,8612,8271,8624143795,0192,758     810800690728204907287283006984154102706055728400     21.14.4206.762.226.713.35.06.713.3-93.3-95.6-90.0-93.3-105.08.962.88.9     70404041448056414456-136-140-130-136-1574814548     36,70736,7076821155085536,01037,75353,34337,75337,7532,98853,34353,34353,3431,6091,380     80,92980,9291,5042531,12012179,39283,235117,60683,23583,2356,589117,606117,606117,6063,5473,042   92.692.63.9711.21025.8194.2794.232694.2394.2399.1126262699.1199.11   3.93.99.5432.57050.633.793.74503.743.740.465050500.460.46     2.482.4885.4453.2510023.560.921.04241.041.040.012424240.010.01   0.980.981.052.97000.980.9500.950.950.280000.280.28   0.040.0400000.040.0400.040.040.140000.140.14
             表6(续)
             能耗
能耗hp 能耗kW
压缩机膨胀过程30的压缩机MCR致冷系统45的压缩机第1级第2级分馏设备35的再压缩机压缩39膨胀机膨胀过程30的膨胀机膨胀机34泵MCR致冷系统45的泵分馏设备35的产品泵净能耗设备总能耗 2,30073,90025,100101,100-2,300-1,10048020 1,71555,10718,7177820-1,715-82035815
99,500106,300 74,19779,268
表7
    相             压力              温度               流量                  组成，摩尔％
   物流 汽(V)/液(L)    kPa     psia     ℃     °F  kg摩尔/hr    lb摩尔/hr    C1    C2    C3+     CO2     N2
    1011121314151617181920212223242526273238414243     V/LV/LLV/LLV/LVVLLLVVLV/LV/LV/LVVVLV/LV/L   5,5855,5164,7575,0191383,3785,0195,0192,6894,6062,8612,6892,8272,2754834142,4823795,0332,7232,482414414     810800690728204907287283906684153954103307060360557303953606060     21.14.4206.762.226.713.35.02.219.4-93.3-95.619.4-90.0-93.3-109.4-51.1-45.612.8-53.312.8-45.6-54.4-51.1     70404041448056413667-136-14067-130-136-165-60-5055-6455-50-66-60     36,70736,7076821155085536,01037,60419,67337,60437,60432,7732,98832,77332,77332,77332,77352,4461,4841,50919,67319,67352,446     80,92980,9291,5042531,12012179,39282,90643,37582,90682,90672,2546,58972,25472,25472,25472,254115,6293,2723,32743,37543,375115,629   92.692.63.9711.21025.8194.2794.226.6694.2294.2235.2199.1135.2135.2135.2135.2124.599.1199.116.666.6624.5   3.93.99.5432.57050.633.793.7545.523.753.7557.810.4657.8157.8157.8157.8153.20.460.4645.5245.5253.2   2.482.4885.4453.2510023.560.921.0347.821.031.036.980.016.986.986.986.9822.30.010.0147.8247.8222.3     0.980.981.052.97000.980.9800.960.960.00.28000000.280.28000     0.040.0400000.040.0400.040.0400.14000000.140.14000
              表7(续)
               能耗
能耗hp 能耗kW
压缩机膨胀过程30的压缩机MCR致冷系统45的压缩机第1级第2级分馏设备35的再压缩机压缩机39膨胀机膨胀过程30的膨胀机膨胀机34泵MCR致冷系统45的泵分馏设备35的产品泵净能耗设备总能耗 2,30080,00031,50010450-2,300-98069020 1,71559,65623,4907336-1,715-73151515
111,700118,300 83,29588,216

Claims (21)

1.一种液化富含甲烷的加压气流的方法，包括以下步骤：在由封闭回路多组分致冷系统冷却的换热器中液化气流，产生温度高于约-112℃(-170°F)并且压力足以使液体产品处于或低于其泡点的富含甲烷的液体产品，将液体产品引入储存装置在高于约-112℃(-170°F)的温度下储存。
2.权利要求1的方法，进一步包括：在将液体产品引入储存装置之前，通过膨胀装置降低液体产品的压力，所述膨胀装置在高于约-112℃(-170°F)的温度和足以使液体产品处于或低于其泡点的压力下产生液体物流。
3.权利要求1的方法，进一步包括：将由于液化天然气蒸发产生的沸腾气体输送到所述换热器，沸腾天然气至少部分地由换热器液化，使液化的沸腾气体加压，所述加压的沸腾气体的温度高于约-112℃(-170°F)，并且压力足以使液体产品处于或低于其泡点。
4.权利要求3的方法，其中换热器包括第一冷却区和操作温度低于第一冷却区的第二冷却区，使权利要求1的气流通过第一冷却区进行液化，使沸腾气体通过第二冷却区进行液化。
5.权利要求4的方法，进一步包括：在沸腾气体通过换热器之前，排出一部分沸腾气体，使排出部分的沸腾气体通过第一冷却区，使排出的沸腾气体升温，冷却换热器中的气流，并将升温的沸腾气体用为燃料。
6.权利要求1的方法，进一步包括：将由于液化天然气蒸发所产生的沸腾气体压缩到被引入换热器的气流的压力，在气流被输送到换热器之前，将压缩的沸腾气体与所述气流合并。
7.权利要求1的方法，进一步包括：使由于液化天然气蒸发所产生的沸腾气体通过换热器以冷却沸腾气体，压缩沸腾气体，使压缩的沸腾气体与气流合并，使合并的沸腾气体与气流通过换热器进行液化。
8.权利要求7的方法，进一步包括：在沸腾气体通过换热器之后，但在压缩冷却的沸腾气体之前，排出一部分沸腾气体，并将排出部分作为燃料。
9.权利要求3的方法，其中换热器包括第一冷却区、第二冷却区和第三冷却区，所述第二冷却区在低于第一冷却区但高于第三冷却区的温度下操作，进一步包括：使沸腾气体通过第三冷却区以液化沸腾气体，在通过第三冷却区之前排出一部分沸腾气体，并使排出的沸腾气体通过第二冷却区，使排出的沸腾气体升温，将升温的排出沸腾气体用作燃料。
10.权利要求1的方法，其中气流含有甲烷和重于甲烷的重质烃，进一步包括通过分馏除去大部分重质烃，以产生富含甲烷的蒸汽物流和富含重质烃的液体物流，然后，蒸汽物流由换热器液化。
11.权利要求10的方法，其中富含重质烃的液体物流被进一步分馏以产生富含乙烷的蒸汽，它与权利要求7的富含甲烷的物流合并。
12.权利要求10的方法，进一步包括在进料物流被分馏之前冷却进料物流。
13.权利要求1的方法，其中换热器包括第一冷却区和第二冷却区，所述第一冷却区是这样被冷却的：使多组分致冷剂液体通过第一冷却区以冷却液体致冷剂，使液体致冷剂经压力膨胀装置进一步降低液体致冷剂的温度，使来自膨胀装置的致冷剂通过第一冷却区，使多组分蒸汽致冷剂通过第一和第二冷却区以降低其温度，使冷却后的蒸汽致冷剂通过膨胀装置，使膨胀的致冷剂通过第二冷却区，然后通过第一冷却区；使气流通过第一冷却区和第二冷却区液化气流，以产生温度高于约-112℃(-170°F)并且压力足以使液体产品处于或低于其泡点的液体产品。
14.权利要求1的方法，进一步包括：
(a)冷却气流以实现气流的部分液化；
(b)将部分冷凝的气流分离成富含重于甲烷的烃的液体物流和富含甲烷的蒸汽物流；
(c)在至少一个分馏塔中将液化部分分馏，以产生富含乙烷的蒸汽物流和富含重于乙烷的烃的液体物流，从过程中除去液体物流；
(d)将富含甲烷的蒸汽物流与富含乙烷的蒸汽物流合并，使合并物流通过权利要求1的换热器，从而液化合并物流；
(e)在将合并液体物流引入储存装置之前，膨胀至少一部分过冷液体以产生高于约-112℃(-170°F)并且压力足以使液体产品处于或低于其泡点的液体产品。
15.权利要求14的方法，其中天然气在步骤(a)中的冷却至少一部分是由封闭回路丙烷致冷系统提供的。
16.权利要求14的方法，进一步包括：使由于液化天然气蒸发所产生的沸腾气体通过换热器，以产生高于约-112℃(-170°F)且压力足以使液体产品处于或低于其泡点的的第二液化天然气，使第二液化天然气物流与权利要求14中步骤(e)的膨胀液化气体合并。
17.权利要求14的方法，其中步骤(d)的换热器包括第一冷却区和操作温度低于第一冷却区的第二冷却区，其中权利要求14的步骤(b)和步骤(c)中的富含甲烷的气流被输送通过第一冷却区进行液化，而由于液化天然气蒸发所产生的温度高于约-112℃(-170°F)的沸腾气体被输送通过第二冷却区进行液化。
18.权利要求10的方法，其中气流在约0℃至约50℃的提高的温度和约2758kPa(400psia)至约8274kPa(1200psia)的提高的压力下进入过程，过程所生产的液化产品的压力高于约1724kPa(250psia)，温度高于约-112℃(-170°F)。
19.权利要求1的方法，其中多组分致冷系统的致冷剂包括甲烷、乙烷、丙烷、丁烷、戊烷、二氧化碳、硫化氢和氮气。
20.一种液化含有甲烷、丙烷、重质烃的天然气流以生产压力高于约1724kPa(250psia)，温度高于约-112℃(-170°F)的液化天然气的方法，该方法包括：
(a)使天然气流通过多组分换热器的第一冷却区，多组分换热器包括三个冷却区，第二冷却区的操作温度低于第一冷却区，但高于第三冷却区；
(b)分馏冷却的天然气流，以分离富含甲烷的物流和重质烃物流；
(c)分馏重质烃物流，以产生富含乙烷的物流和含有重于乙烷的烃的物流，从过程中除去重于乙烷的烃；
(d)将步骤(b)的富含甲烷的物流与步骤(c)的富含乙烷的物流合并，使合并物流通过多组分致冷系统的第二冷却区，冷却合并的物流以产生过冷的冷凝物；
(e)使至少一部分过冷的冷凝物膨胀以提供压力高于约1724kPa(250psia)，温度高于约-112℃(-170°F)的液化天然气；
(f)使由于储存装置内的液化天然气蒸发所产生的气体通过多组分致冷系统的第三冷却区，以产生第二液化天然气物流，使第二液化天然气物流与步骤(e)中产生的液化天然气物流合并。
21.一种液化含有甲烷、丙烷、重质烃的天然气流以生产压力高于约1724kPa(250psia)，温度高于约-112℃(-170°F)的液化天然气的方法，该方法包括：
(a)通过丙烷致冷系统冷却天然气流；
(b)分馏冷却的天然气流，以分离富含甲烷的物流和重质烃物流；
(c)分馏重质烃物流，以产生富含乙烷的物流和至少一股含有重于乙烷的烃的物流，从过程中除去重于乙烷的烃；
(d)将步骤(b)的富含甲烷的物流与步骤(c)的富含乙烷的物流合并，使合并物流通过多组分致冷系统的第一冷却区，致冷系统的第一冷却区由多组分液体和多组分蒸汽在换热器中冷却，所述的多组分液体和多组分蒸汽与合并的富含甲烷物流和富含乙烷物流换热以产生过冷冷凝物；和
(e)使至少一部分过冷的冷凝物膨胀以提供压力高于约1724kPa(250psia)，温度高于约-112℃(-170°F)的液化天然气；
(f)使由于储存装置内的液化天然气蒸发所产生的气体通过多组分致冷系统的第三冷却区，以产生第二液化天然气物流，使第二液化天然气物流与步骤(e)中产生的液化天然气物流合并。
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