WO1997036139A1 - Aromatics and/or heavies removal from a methane-based feed by condensation and stripping - Google Patents

Abstract

A method and associated apparatus for removal of benzene, other aromatics and/or other heavier hydrocarbon components from a methane-based gas stream by condensation and stripping. It is desirable to remove benzene and other aromatics to prevent fouling and plugging of processing equipment and it is desirable to recover other heavier hydrocarbon components because of their value. Cooled feed stream (118) is fed to a column (60) and separated into methane-rich vapor stream (120) and benzene/aromatics/heavies liquid (114). The liquid (114) is sent to a heat exchanger (62) to recover refrigeration. Warm dry gas (108) is cooled in the heat exchanger (62) and delivered as stripping gas (109) to the column (60).

Description

AROMATICS AND/OR HEAVIES REMOVAL FROM A METHANE-BASED FEED BY CONDENSATION AND STRIPPING

This invention concerns a method and associated apparatus for removing benzene,

other aromatics and/or heavier hydrocarbon components from a methane-based gas stream by a

unique condensation and stripping process.

BACKGROUND

Cryogenic liquefaction of normally gaseous materials is utilized for the purposes

of component separation, purification, storage and for the transportation of said components in a

more economic and convenient form. Most such liquefaction systems have many operations in

common, regardless ofthe gases involved, and consequently, have many ofthe same problems.

One problem commonly encountered in liquefaction processes, particularly when aromatics are

present, is the precipitation and subsequent solidification of these species in the process

equipment thereby resulting in reduced process efficiency and reliability. Another common

problem is the removal of small quantities ofthe higher valued, higher molecular weight

chemical species from the gas stream immediately prior to liquefaction ofthe gas stream in a

major portion. Accordingly, the present invention will be described with specific reference to the

processing of natural gas but is applicable to the processing of gas in other systems wherein

similar problems are encountered. It is common practice in the art of processing natural gas to subject the gas to

cryogenic treatment to separate hydrocarbons having a molecular weight higher than methane

(C₂+) from the natural gas thereby producing a pipeline gas predominating in methane and a C₂+

stream useful for other purposes. Frequently, the C₂+ stream will be separated into individual

component streams, for example, C₂, C₃, C₄ and C_$+.

It is also common practice to cryogenically treat natural gas to liquefy the same

for transport and storage. The primary reason for the liquefaction of natural gas is that

liquefaction results in a volume reduction of about 1/600, thereby making it possible to store and

transport the liquefied gas in containers of more economical and practical design. For example,

when gas is transported by pipeline from the source of supply to a distant market, it is desirable

to operate the pipeline under a substantially constant and high load factor. Often the

deliverability or capacity ofthe pipeline will exceed demand while at other times the demand

may exceed the deliverability ofthe pipeline. In order to shave off the peaks where demand

exceeds supply, it is desirable to store the excess gas in such a manner that it can be delivered

when the supply exceeds demand, thereby enabling future peaks in demand to be met with

material from storage. One practical means for doing this is to convert the gas to a liquefied state

for storage and to then vaporize the liquid as demand requires.

Liquefaction of natural gas is of even greater importance in making possible the

transport of gas from a supply source to market when the source and market are separated by

great distances and a pipeline is not available or is not practical. This is particularly true where

transport must be made by ocean-going vessels. Ship transportation in the gaseous state is

generally not practical because appreciable pressurization is required to significantly reduce the

specific volume ofthe gas which in turn requires the use of more expensive storage containers. In order to store and transport natural gas in the liquid state, the natural gas is

preferably cooled to -240°F to -260 °F where it possesses a near-atmospheric vapor pressure.

Numerous systems exist in the prior art for the liquefaction of natural gas or the like in which the

gas is liquefied by sequentially passing the gas at an elevated pressure through a plurality of

cooling stages whereupon the gas is cooled to successively lower temperatures until the

liquefaction temperature is reached. Cooling is generally accomplished by heat exchange with

one or more refrigerants such as propane, propylene, ethane, ethylene, and methane or a

combination of one or more ofthe preceding. In the art, the refrigerants are frequently arranged

in a cascaded manner and each refrigerant is employed in a closed refrigeration cycle. Further

cooling ofthe liquid is possible by expanding the liquefied natural gas to atmospheric pressure in

one or more expansion stages. In each stage, the liquefied gas is flashed to a lower pressure

thereby producing a two-phase gas-liquid mixture at a significantly lower temperature. The

liquid is recovered and may again be flashed. In this manner, the liquefied gas is further cooled

to a storage or transport temperature suitable for liquefied gas storage at near-atmospheric

pressure. In this expansion to near-atmospheric pressure, some additional volumes of liquefied

gas are flashed. The flashed vapors from the expansion stages are generally collected and

recycled for liquefaction or utilized as fuel gas for power generation.

As previously noted, a major operational problem in the liquefaction of natural

gas is the removal of residual amounts of benzene and other aromatic compounds from the

natural gas stream immediately prior to the liquefaction ofa major portion of said stream and the

tendency of such components to precipitate and solidify thereby causing the fouling and potential

plugging of pipes and key process equipment. As an example, such fouling can significantly reduce the heat transfer efficiency and throughput of heat exchangers, particularly plate-fin heat exchangers.

For technical and economic reasons it is not necessary to remove impurities such

as benzene completely. It is, however, desirable to reduce its concentration. Contaminant

removal from natural gas may be accomplished by the same type of cooling used in the

liquefaction process wherein the contaminants condense in accordance with their respective

condensation temperature. Except for the fact that the gas must be cooled to a lower temperature to liquefy, as opposed to separating the benzene contaminant, the basic cooling techniques are

the same for liquefaction and separation. Accordingly, in respect of residual benzene, it is only

necessary to cool the natural gas to a temperature at which a portion ofthe feed gas is condensed. This may be accomplished in a cryogenic separation column included at an appropriate point in

the LNG recovery process to separate the condensed benzene from the main gas stream.

In the interest of efficient operation ofthe cryogenic separation column, it is

desirable to utilize the condensed liquid at cryogenic temperatures, that must be withdrawn from

the column, for heat exchange with a warm dry gas stream provided to the cryogenic separation

column. This heat exchange scheme, however, presents a problem resulting from the excessive

temperature differential ofthe two streams supplied to the heat exchanger. Since the actual

temperature difference could exceed 100°F, the thermal shock to the heat exchanger could

damage or shorten useful life ofthe heat exchanger apparatus constructed of conventional

materials.

Another consideration related to efficient operation of a cryogenic separation

column is providing heat exchanger controls that allow automatic start-up ofthe column. Still yet another problem in the processing of methane-rich gas streams is the lack

ofa cost-effective means for recovering the higher molecular weight hydrocarbons from the gas

stream prior to liquefaction ofthe stream in major portion or returning the remaining stream to a

pipeline or other processing step. The recovered higher molecular weight hydrocarbons

generally possess a greater value on a per unit mass basis than the remaining components in the

gas stream.

SUMMARY OF THE INVENTION

It is an object of this invention to remove residual quantities of benzene and other

aromatics from a methane-based gas stream which is to be liquefied in major portion.

It is another object of this invention to remove the higher molecular weight

hydrocarbons from a methane-based gas stream.

It is still yet another object of this invention to remove the higher molecular

weight hydrocarbons from a methane-based gas stream which is to be liquefied in a major

portion.

It is yet still further an object of this invention to remove benzene, other aromatics

and/or the higher molecular weight hydrocarbons from methane-based gas stream in an energy-

efficient manner.

It is still further an object ofthe present invention that the process employed for

the removal of benzene, other aromatics and/or higher molecular weight hydrocarbons be

compatible with and integrate into technology routinely employed in gas plants.

And further yet still, it is an object of this invention that the process and apparatus

employed for benzene, other aromatic and/or high molecular weight hydrocarbon removal from

a methane-based gas stream be relatively simple, compact and cost-effective. It still further yet is an object of the present invention that the process employed

for the removal of benzene, other aromatics and/or higher molecular hydrocarbons from a

methane-based gas stream to be liquefied in major portion be compatible with and integrate into

technology routinely employed in plants producing liquefied natural gas.

Yet still further an object of this is invention is to provide heat exchanger controls

which overcome the above-mentioned and other associated problems in handling low

temperature fluids.

Another object of this invention is to provide an improved control method which reduces initial equipment temperature requirements, and costs for heat exchange apparatus.

A more specific object is to control heat exchanger temperatures to allow cooling

of a warm fluid stream against a low temperature fluid stream without introducing thermal shock

to the heat exchange apparatus.

A still further object of this invention is to control the heat exchanger to facilitate

automatic start-up of a cryogenic separation column.

In one embodiment of this invention, benzene and/or other aromatics are removed

from a methane-based gas stream by a process comprising (1) condensing a minor portion ofthe

methane-based gas stream immediately prior to the step wherein a majority of said gas stream is

liquefied thereby producing a two-phase stream, (2) feeding said two-phase stream into the upper

section of a stripping column, (3) removing from the upper section of said stripping column an

aromatic-depleted gas stream, (4) removing from the lower section of said stripping column an

aromatic-rich liquid stream, (5) contacting via indirect heat exchange the aromatic-rich liquid

stream with a methane-rich stripping gas stream thereby producing a warmed aromatic-bearing stream and a cooled methane-rich stripping gas stream, and (6) feeding said cooled methane-rich stripping gas stream to the lower section ofthe stripping column, and optionally (7) feeding said

aromatic-depleted gas stream to a liquefaction step wherein the gas stream is liquefied in major

portion thereby producing liquefied natural gas.

In another embodiment of this invention, the higher molecular weight

hydrocarbons in a methane-based gas stream are removed and concentrated by a process

comprising (1) condensing a minor portion ofthe methane-based gas stream to produce a two-

phase stream, (2) feeding said two-phase stream into the upper section of a stripping column, (3)

removing from the upper section of said stripping column a heavies-depleted gas stream, (4)

removing the lower section of said stripping column a heavies-rich liquid stream, (5) contacting

via indirect heat exchange the heavies-rich liquid stream with a methane-rich stripping gas

stream thereby producing a warmed heavies-rich stream and a cooled methane-rich stripping gas

stream, and (6) feeding said methane-rich stripping gas stream to the lower section ofthe

stripping column.

In still yet another embodiment of this invention, the invention is an apparatus

comprising (1) a condenser wherein a minor portion ofa methane-based gas stream is condensed

thereby producing a two-phase stream, (2) a stripping column to which the two-phase stream is

fed and from which is produced a vapor stream and a liquid stream, (3) a heat exchanger

containing an indirect heat exchange means which provides for indirect heat exchange between a

gas stream and the liquid stream thereby producing a cooled gas stream and a warmed liquid

stream, (4) a conduit between said condenser and the upper section ofthe stripping column for

flow of said two-phase stream, (5) a conduit connected to the upper section of the stripping

column for removal of said vapor stream, (6) a conduit between said stripping column and said

heat exchanger for flow of said liquid stream, (7) a conduit between said heat exchanger and said stripping column for flow of said cooled gas stream, (8) a conduit connected to said heat

exchanger for the flow of a said warmed liquid stream from the heat exchanger, and (9) a

conduit connected to said heat exchanger for flow of said gas stream to the heat exchanger.

In yet another embodiment of this invention, the foregoing and other objectives

and advantages are achieved in controlling a heat exchanger handling a low temperature fluid

and a warm fluid by providing a by-pass conduit for the warm fluid, wherein a control valve in

the by-pass conduit is manipulated responsive to the temperature ratio ofthe heat exchange

fluids. In accordance with another aspect ofthe invention automatic start-up controls include a

high selector for temporarily selecting a temperature to manipulate flow ofthe warm fluid that

facilitates start-up of the column, and then switches to manipulation of the warm gas flow

responsive to a desired temperature.

BRIEF DESCRIPTION OF THE DRAWINGS

FIGURE 1 is a simplified flow diagram of a cryogenic LNG production process

which illustrates the methodology and apparatus ofthe present invention for the removal of

benzene, other aromatics and/or higher molecular weight hydrocarbon species from a methane-

based gas stream.

FIGURE 2 is a simplified flow diagram which illustrates in greater detail the

methodology and apparatus illustrated in FIGURE 1.

FIGURE 3 is a diagrammatic illustration of a cryogenic separation column and

the associated control system ofthe present invention for maintaining a desired temperature ratio

for the heat exchange fluids.

FIGURE 4 is a diagrammatic illustration similar to FIGURE 3 for temporarily

selecting a temperature that will allow automatic start-up ofthe cryogenic separation column. DESCRIPTION OF THE PREFERRED EMBODIMENTS

While the present invention in the preferred embodiments is applicable to (1 ) the

removal of benzene and/or other aromatics from a methane-based gas stream which is to be

condensed in major portion and (2) the removal ofthe more valuable, higher molecular weight

hydrocarbon species from a methane-based gas stream which is to be condensed in major

portion, the technology is also applicable to the generic recovery of such species from methane-

based streams (e.g., removal of natural gas liquids from natural gas). Benzene and other

aromatics present a unique problem because of their relatively high melting point temperatures.

As an example, benzene which contains 6 carbon atoms possesses a melting point of 5.5 °C and a

boiling point of 80.1 °C. Hexane, which also contains 6 carbon atoms, possesses a melting point

of -95 °C and a boiling point of 68.95 °C. Therefore when compared to other hydrocarbons of

similar molecular weight, benzene and other aromatic compounds pose a much greater problem

with regard to fouling and/or plugging of process equipment and conduit. Aromatic compounds

as used herein are those compounds characterized by the presence of at least one benzene ring.

As used herein, higher molecular hydrocarbon species are those hydrocarbon species possessing,

molecular weight greater than ethane, and this term will be used interchangeably with heavy

hydrocarbons.

For the purposes of simplicity and clarity, the following description will be

confined to the employment ofthe inventive processes and associated apparatus in the cryogenic

cooling of a natural gas stream to produce liquefied natural gas. More specifically, the following

description will focus on the removal of benzene and/or other aromatic species and/or higher

molecular weight hydrocarbons (heavy hydrocarbons) in a liquefaction scheme wherein cascaded

refrigeration cycles are employed. However, the applicability ofthe inventive processes and associated apparatus herein described is not limited to liquefaction systems which employ

cascaded refrigeration cycles or which process natural gas streams exclusively. The processes

and associated apparatus are applicable to any refrigeration system wherein (a) benzene and/or

heavier aromatics exist in a methane-based gas stream at concentrations which may foul or plug

process equipment, particularly the heat exchangers employed for condensing said stream, or

(b) it is desirable for whatever reason to remove and recover higher molecular weight

hydrocarbons from a methane-based gas stream.

Natural Gas Stream Liquefaction

Cryogenic plants have a variety of forms; the most efficient and effective being a

cascade-type operation and this type in combination with expansion-type cooling. Also, since

methods for the production of liquefied natural gas (LNG) include the separation of

hydrocarbons of molecular weight greater than methane as a first part thereof, a description of a

plant for the cryogenic production of LNG effectively describes a similar plant for removing C₂+

hydrocarbons from a natural gas stream.

In the preferred embodiment which employs a cascaded refrigerant system, the

invention concerns the sequential cooling of a natural gas stream at an elevated pressure, for

example about 650 psia, by sequentially cooling the gas stream by passage through a multistage

propane cycle, a multistage ethane or ethylene cycle and either (a) a closed methane cycle

followed by a single- or a multistage expansion cycle to further cool the same and reduce the

pressure to near-atmospheric or (b) an open-end methane cycle which utilizes a portion of the

feed gas as a source of methane and which includes therein a multistage expansion cycle to

further cool the same and reduce the pressure to near-atmospheric pressure. In the sequence of cooling cycles, the refrigerant having the highest boiling point is utilized first followed by a

refrigerant having an intermediate boiling point and finally by a refrigerant having the lowest

boiling point.

Pretreatment steps provide a means for removing undesirable components such as

acid gases, mercaptans, mercury and moisture from the natural gas feed stream delivered to the

facility. The composition of this gas stream may vary significantly. As used herein, a natural

gas stream is any stream principally comprised of methane which originates in major portion

from a natural gas feed stream, such feed stream for example containing at least 85% methane by

volume, with the balance being ethane, higher hydrocarbons, nitrogen, carbon dioxide and a

minor amounts of other contaminants such as mercury, hydrogen sulfide, mercaptans. The

pretreatment steps may be separate steps located either upstream ofthe cooling cycles or located

downstream of one ofthe early stages of cooling in the initial cycle. The following is a non-

inclusive listing of some ofthe available means which are readily available to one skilled in the

art. Acid gases and to a lesser extent mercaptans are routinely removed via a sorption process

employing an aqueous amine-bearing solution. This treatment step is generally performed

upstream ofthe cooling stages employed in the initial cycle. A major portion ofthe water is

routinely removed as a liquid via two-phase gas-liquid separation following gas compression and

cooling upstream ofthe initial cooling cycle and also downstream ofthe first cooling stage in the

initial cooling cycle. Mercury is routinely removed via mercury sorbent beds. Residual amounts

of water and acid gases are routinely removed via the use of properly selected sorbent beds such

as regenerable molecular sieves. Processes employing sorbent beds are generally located

downstream ofthe first cooling stage in the initial cooling cycle. The resulting natural gas stream is generally delivered to the liquefaction process

at an elevated pressure or is compressed to an elevated pressure, that being a pressure greater

than 500 psia, preferably about 500 to about 900 psia, still more preferably about 550 to about

675 psia, still yet more preferably about 575 to about 650 psia, and most preferably about 600

psia. The stream temperature is typically near ambient to slightly above ambient. A

representative temperature range being 60 °F to 120°F.

As previously noted, the natural gas stream at this point is cooled in a plurality of

multistage (for example, three) cycles or steps by indirect heat exchange with a plurality of

refrigerants, preferably three. The overall cooling efficiency for a given cycle improves as the

number of stages increases but this increase in efficiency is accompanied by corresponding

increases in net capital cost and process complexity. The feed gas is preferably passed through

an effective number of refrigeration stages, nominally two, preferably two to four, and more

preferably three stages, in the first closed refrigeration cycle utilizing a relatively high boiling

refrigerant. Such refrigerant is preferably comprised in major portion of propane, propylene or

mixtures thereof, more preferably propane, and most preferably the refrigerant consists

essentially of propane. Thereafter, the processed feed gas flows through an effective number of

stages, nominally two, preferably two to four, and more preferably two or three, in a second

closed refrigeration cycle in heat exchange with a refrigerant having a lower boiling point. Such

refrigerant is preferably comprised in major portion of ethane, ethylene or mixtures thereof, more

preferably ethylene, and most preferably the refrigerant consists essentially of ethylene. Each of

the above-cited cooling stages for each refrigerant comprises a separate cooling zone.

Generally, the natural gas feed stream will contain such quantities of C₂+

components so as to result in the formation of a C₂+ rich liquid in one or more of the cooling stages. This liquid is removed via gas-liquid separation means, preferably one or more

conventional gas-liquid separators. Generally, the sequential cooling ofthe natural gas in each

stage is controlled so as to remove as much as possible ofthe C₂ and higher molecular weight

hydrocarbons from the gas to produce a first gas stream predominating in methane and a second

liquid stream containing significant amounts of ethane and heavier components. An effective

number of gas/liquid separation means are located at strategic locations downstream of the

cooling zones for the removal of liquids streams rich in C₂+ components. The exact locations

and number of gas/liquid separators will be dependant on a number of operating parameters, such

as the C₂+ composition ofthe natural gas feed stream, the desired BTU content ofthe final

product, the value ofthe C₂+ components for other applications and other factors routinely

considered by those skilled in the art of LNG plant and gas plant operation. The C₂+

hydrocarbon stream or streams may be demethanized via a single stage flash or a fractionation

column. In the former case, the methane-rich stream can be repressurized and recycled or can be

used as fuel gas. In the latter case, the methane-rich stream can be directly returned at pressure

to the liquefaction process. The C₂+ hydrocarbon stream or streams or the demethanized C +

hydrocarbon stream may be used as fuel or may be further processed such as by fractionation in

one or more fractionation zones to produce individual streams rich in specific chemical

constituents (ex., C₂, C₃, C₄ and C₅+). In the last stage ofthe second cooling cycle, the gas

stream which is predominantly methane is condensed (i.e., liquefied) in major portion, preferably

in its entirety. In one ofthe preferred embodiments to be discussed in greater detail in a later

section, it is at this location in the process that the inventive process and associated apparatus for

benzene, other aromatics and/or heavier hydrocarbon removal can be employed. The process pressure at this location is only slightly lower than the pressure of the feed gas to the first stage ofthe first cycle.

The liquefied natural gas stream is then further cooled in a third step or cycle by

one of two embodiments. In one embodiment, the liquefied natural gas stream is further cooled

by indirect heat exchange with a third closed refrigeration cycle wherein the condensed gas

stream is subcooled via passage through an effective number of stages, nominally 2; preferably 2

to 4; and most preferably 3 wherein cooling is provided via a third refrigerant having a boiling

point lower than the refrigerant employed in the second cycle. This refrigerant is preferably

comprised in major portion of methane and more preferably is predominantly methane. In the

second and preferred embodiment which employs an open methane refrigeration cycle, the

liquefied natural gas stream is subcooled via contact with flash gases in a main methane

economizer in a manner to be described later.

In the fourth cycle or step, the liquefied gas is further cooled by expansion and

separation ofthe flash gas from the cooled liquid. In a manner to be described, nitrogen removal

from the system and the condensed product is accomplished either as part of this step or in a

separate succeeding step. A key factor distinguishing the closed cycle from the open cycle is the

initial temperature of the liquefied stream prior to flashing to near-atmospheric pressure, the

relative amounts of flashed vapor generated upon said flashing, and the disposition ofthe flashed

vapors. Whereas the majority ofthe flash vapor is recycled to the methane compressors in the

open-cycle system, the flashed vapor in a closed-cycle system is generally utilized as a fuel.

In the fourth cycle or step in either the open- or closed-cycle methane systems, the

liquefied product is cooled via at least one, preferably two to four, and more preferably three

expansions where each expansion employs either Joule-Thomson expansion valves or hydraulic expanders followed by a separation ofthe gas-liquid product with a separator. When a hydraulic

expander is employed and properly operated, the greater efficiencies associated with the recovery

of power, a greater reduction in stream temperature, and the production of less vapor during the

flash step will frequently be cost-effective even in light of increased capital and operating costs

associated with the expander. In one embodiment employed in the open-cycle system,

additional cooling ofthe high pressure liquefied product prior to flashing is made possible by

first flashing a portion of this stream via one or more hydraulic expanders and then via indirect

heat exchange means employing said flashed stream to cool the high pressure liquefied stream

prior to flashing. The flashed product is then recycled via return to an appropriate location,

based on temperature and pressure considerations, in the open methane cycle.

When the liquid product entering the fourth cycle is at the preferred pressure of

about 600 psia, representative flash pressures for a three stage flash process are about 190, 61

and 14.7 psia. In the open-cycle system, vapor flashed or fractionated in the nitrogen separation

step to be described and that flashed in the expansion flash steps are utilized as cooling agents in

the third step or cycle which was previously mentioned. In the closed-cycle system, the vapor

from the flash stages may also be employed as a cooling agent prior to either recycle or use as

fuel. In either the open- or closed-cycle system, flashing ofthe liquefied stream to near

atmospheric pressure will produce an LNG product possessing a temperature of

-240°F to -260°F.

To maintain the BTU content ofthe liquefied product at an acceptable limit when

appreciable nitrogen exists in the feed stream, nitrogen must be concentrated and removed at

some location in the process. Various techniques for this purpose are available to those skilled in

the art. The following are examples. When an open methane cycle is employed and nitrogen concentration in the feed is low, typically less than about 1.0 vol%, nitrogen removal is generally

achieved by removing a small side stream at the high pressure inlet or outlet port at the methane compressor. For a closed cycle at nitrogen concentrations of up to 1.5 vol.% in the feed gas, the

liquefied stream is generally flashed from process conditions to near-atmospheric pressure in a

single step, usually via a flash drum. The nitrogen-bearing flash vapors are then generally

employed as fuel gas for the gas turbines which drive the compressors. The LNG product which is now at near-atmospheric pressure is routed to storage. When the nitrogen concentration in the inlet feed gas is about 1.0 to about 1.5 vol% and an open-cycle is employed, nitrogen can be

removed by subjecting the liquefied gas stream from the third cooling cycle to a flash step prior

to the fourth cooling step. The flashed vapor will contain an appreciable concentration of nitrogen and may be subsequently employed as a fuel gas. A typical flash pressure for nitrogen removal at these concentrations is about 400 psia. When the feed stream contains a nitrogen

concentration of greater than about 1.5 vol% and an open or closed cycle is employed, the flash step may not provide sufficient mtrogen removal. In such event, a nitrogen rejection column will

be employed from which is produced a nitrogen rich vapor stream and a liquid stream. In a preferred embodiment which employs a nitrogen rejection column, the high pressure liquefied

methane stream to the methane economizer is split into a first and second portion. The first

portion is flashed to approximately 400 psia and the two-phase mixture is fed as a feed stream to

the nitrogen rejection column. The second portion ofthe high pressure liquefied methane stream

is further cooled by flowing through a methane economizer to be described later, it is then flashed to 400 psia, and the resulting two-phase mixture or the liquid portion thereof is fed to the upper section ofthe column where it functions as a reflux stream reflux. The nitrogen-rich vapor

stream produced from the top ofthe nitrogen rejection column will generally be used as fuel. The liquid stream produced from the bottom of the column is then fed to the first stage of

methane expansion.

Refrigerative Cooling for Natural Gas Liquefaction

Critical to the liquefaction of natural gas in a cascaded process is the use of one or

more refrigerants for transferring heat energy from the natural gas stream to the refrigerant and

ultimately transferring said heat energy to the environment. In essence, the refrigeration system

functions as a heat pump by removing heat energy from the natural gas stream as the stream is

progressively cooled to lower and lower temperatures.

The liquefaction process employs several types of cooling which include but are

not limited to (a) indirect heat exchange, (b) vaporization and (c) expansion or pressure

reduction. Indirect heat exchange, as used herein, refers to a process wherein the refrigerant or

cooling agent cools the substance to be cooled without actual physical contact between the

refrigerating agent and the substance to be cooled. Specific examples include heat exchange

undergone in a tube-and-shell heat exchanger, a core-in-kettle heat exchanger, and a brazed

aluminum plate-fin heat exchanger. The physical state of the refrigerant and substance to be

cooled can vary depending on the demands ofthe system and the type of heat exchanger chosen.

Thus, in the inventive process, a shell-and-tube heat exchange will typically be utilized where the

refrigerating agent is in a liquid state and the substance to be cooled is in a liquid or gaseous

state, whereas, a plate-fin heat exchanger will typically be utilized where the refrigerant is in a

gaseous state and the substance to be cooled is in a liquid state. Finally, the core-in-kettle heat

exchanger will typically be utilized where the substance to be cooled is liquid or gas and the refrigerant undergoes a phase change from a liquid state to a gaseous state during the heat exchange.

Vaporization cooling refers to the cooling of a substance by the evaporation or

vaporization of a portion ofthe substance with the system maintained at a constant pressure.

Thus, during the vaporization, the portion ofthe substance which evaporates absorbs heat from

the portion ofthe substance which remains in a liquid state and hence, cools the liquid portion.

Finally, expansion or pressure reduction cooling refers to cooling which occurs

when the pressure of a gas-, liquid- or a two-phase system is decreased by passing through a

pressure reduction means. In one embodiment, this expansion means is a Joule-Thomson

expansion valve. In another embodiment, the expansion means is a hydraulic or gas expander.

Because expanders recover work energy from the expansion process, lower process stream

temperatures are possible upon expansion.

In the discussion and drawings to follow, the discussions or drawings may depict

the expansion of a refrigerant by flowing through a throttle valve followed by a subsequent

separation of gas and liquid portions in the refrigerant chillers or condensers, as the case may be,

wherein indirect heat-exchange also occurs. While this simplified scheme is workable and

sometimes preferred because of cost and simplicity, it may be more effective to carry out

expansion and separation and then partial evaporation as separate steps, for example a

combination of throttle valves and flash drums prior to indirect heat exchange in the chillers or

condensers. In another workable embodiment, the throttle or expansion valve may not be a

separate item but an integral part of the refrigerant chiller or condenser (i.e., the flash occurs

upon entry ofthe liquefied refrigerant into the chiller). In a like manner, the cooling of multiple streams for a given refrigeration stage may occur within a single vessel (i.e., chiller) or within

multiple vessels. The former is generally preferred from a capital equipment cost perspective.

In the first cooling cycle, cooling is provided by the compression of a higher

boiling point gaseous refrigerant, preferably propane, to a pressure where it can be liquefied by

indirect heat transfer with a heat transfer medium which ultimately employs the environment as a

heat sink, that heat sink generally being the atmosphere, a fresh water source, a salt water source,

the earth or two or more ofthe preceding. The condensed refrigerant then undergoes one or more

steps of expansion cooling via suitable expansion means thereby producing two-phase mixtures

possessing significantly lower temperatures. In one embodiment, the main stream is split into at

least two separate streams, preferably two to four streams, and most preferably three streams

where each stream is separately expanded to a designated pressure. Each stream then provides

evaporative cooling via indirect heat transfer with one or more selected streams, one such stream

being the natural gas stream to be liquefied. The number of separate refrigerant streams will

correspond to the number of refrigerant compressor stages. The vaporized refrigerant from each

respective stream is then returned to the appropriate stage at the refrigerant compressor (e.g., two

separate streams will correspond to a two-stage compressor). In a more preferred embodiment,

all liquefied refrigerant is expanded to a predesignated pressure and this stream then employed to

provide vaporative cooling via indirect heat transfer with one or more selected streams, one such

stream being the natural gas stream to be liquefied. A portion ofthe liquefied refrigerant is then

removed from the indirect heat transfer means, expansion cooled by expanding to a lower

pressure and correspondingly lower temperature where it provides vaporative cooling via indirect

heat transfer means with one or more designated streams, one such stream being the natural gas

stream to be liquefied. Nominally, this embodiment will employ two such expansion cooling/vaporative cooling steps, preferably two to four, and most preferably three. Like the first

embodiment, the refrigerant vapor from each step is returned to the appropriate inlet port at the staged compressor.

In the preferred cascaded embodiment, the majority ofthe cooling for liquefaction

ofthe lower boiling point refrigerants (i.e., the refrigerants employed in the second and third

cycles) is made possible by cooling these streams via indirect heat exchange with selected

higher boiling refrigerant streams. This manner of cooling is referred to as "cascaded cooling."

In effect, the higher boiling refrigerants function as heat sinks for the lower boiling refrigerants

or stated differently, heat energy is pumped from the natural gas stream to be liquefied to a lower

boiling refrigerant and is then pumped (i.e., transferred) to one or more higher boiling

refrigerants prior to transfer to the environment via an environmental heat sink (ex., fresh water,

salt water, atmosphere). As in the first cycle, refrigerant employed in the second and third

cycles are compressed via multi-staged compressors to preselected pressures. When possible and

economically feasible, the compressed refrigerant vapor is first cooled via indirect heat exchange

with one or more cooling agents (ex., air, salt water, fresh water) directly coupled to environmental heat sinks. This cooling may be via inter-stage cooling between compression

stages and/or cooling ofthe compressed product. The compressed stream is then further cooled

via indirect heat exchange with one or more ofthe previously discussed cooling stages for the

higher boiling point refrigerants. The second cycle refrigerant, preferably ethylene, is preferably first cooled via

indirect heat exchange with one or more cooling agents directly coupled to an environmental heat

sink (i.e., inter-stage and/or post-cooling following compression) and then further cooled and

finally liquefied via sequential contact with the first and second or first, second and third cooling stages for the highest boiling point refrigerant which is employed in the first cycle. The

preferred second and first cycle refrigerants are ethylene and propane, respectively.

When employing a three refrigerant cascaded closed cycle system, the refrigerant

in the third cycle is compressed in a stagewise manner, preferably though optionally cooled via

indirect heat transfer to an environmental heat sink (i.e., inter-stage and/or post-cooling

following compression) and then cooled by indirect heat exchange with either all or selected

cooling stages in the first and second cooling cycles which preferably employ propane and

ethylene as respective refrigerants. Preferably, this stream is contacted in a sequential manner

with each progressively colder stage of refrigeration in the first and second cooling cycles,

respectively.

In an open-cycle cascaded refrigeration system such as that illustrated in FIGURE

1 , the first and second cycles are operated in a manner analogous to that set forth for the closed

cycle. However, the open methane cycle system is readily distinguished from the conventional

closed refrigeration cycles. As previously noted in the discussion ofthe fourth cycle or step, a

significant portion ofthe liquefied natural gas stream originally present at elevated pressure is

cooled to approximately -260 °F by expansion cooling in a stepwise manner to near-atmospheric

pressure. In each step, significant quantities of methane-rich vapor at a given pressure are

produced. Each vapor stream preferably undergoes significant heat transfer in methane

economizers and is preferably returned to the inlet port ofa compressor stage at near-ambient

temperature. In the course of flowing through the methane economizers, the flashed vapors are

contacted with warmer streams in a countercurrent manner and in a sequence designed to

maximize the cooling ofthe warmer streams. The pressure selected for each stage of expansion

cooling is such that for each stage, the volume of gas generated plus the compressed volume of vapor from the adjacent lower stage results in efficient overall operation of the multi-staged

compressor. Interstage cooling and cooling ofthe final compressed gas is preferred and

preferably accomplished via indirect heat exchange with one or more cooling agents directly

coupled to an environmental heat sink. The compressed methane-rich stream is then further

cooled via indirect heat exchange with refrigerant in the first and second cycles, preferably all

stages associated with the refrigerant employed in the first cycle, more preferably the first two

stages and most preferably, only the first stage. The cooled methane-rich stream is further

cooled via indirect heat exchange with flash vapors in the main methane economizer and is then

combined with the natural gas feed stream at a location in the liquefaction process where the

natural gas feed stream and the cooled methane-rich stream are at similar conditions of

temperature and pressure, preferably prior to entry into one of the stages of ethylene cooling,

more preferably immediately prior to the ethylene cooling stage wherein methane in major

portion is liquefied (i.e., ethylene condenser).

Optimization via Inter-stage and Inter-cycle Heat Transfer

In the more preferred embodiments, steps are taken to further optimize process

efficiency by returning the refrigerant gas streams to the inlet port of their respective

compressors at or near ambient temperature. Not only does this step improve overall

efficiencies, but difficulties associated with the exposure of compressor components to cryogenic

conditions are greatly reduced. This is accomplished via the use of economizers wherein streams

comprised in major portion of liquid and prior to flashing are first cooled by indirect heat

exchange with one or more vapor streams generated in a downstream expansion step (i.e., stage)

or steps in the same or a downstream cycle. In a closed system, economizers are preferably employed to obtain additional cooling from the flashed vapors in the second and third cycles.

When an open methane cycle system is employed, flashed vapors from the fourth stage are

preferably returned to one or more economizers where (1) these vapors cool via indirect heat

exchange the liquefied product streams prior to each pressure reduction stage and (2) these

vapors cool via indirect heat exchange the compressed vapors from the open methane cycle prior

to combination of this stream or streams with the main natural gas feed stream. These cooling

steps comprise the previously discussed third stage of cooling and will be discussed in greater

detail in the discussion of FIGURE 1. In one embodiment wherein ethylene and methane are

employed in the second and third cycles, the contacting can be performed via a series of ethylene

and methane economizers. In a preferred embodiment which is illustrated in FIGURE 1 and which will be discussed in greater detail later, the process employs a main ethylene economizer,

a main methane economizer and one or more additional methane economizers. These additional

economizers are referred to herein as the second methane economizer, the third methane

economizer and so forth and each such additional methane economizer corresponds to a separate

downstream flash step.

Benzene. Other Aromatic and/or Heavier Hydrocarbon Removal

The inventive process for the removal of benzene, other aromatics and/or the

higher molecular weight hydrocarbon species from a methane-based gas stream is an extremely

energy efficient and operationally simple process. Because ofthe manner of operation, the

column referred to herein as a stripping column performs both stripping and fractionating

functions. The process comprises cooling the methane-based gas stream such that 0.1 to 20

mol%, preferably 0.5 to about 10 mol%, and more preferably about 1.75 to about 6.0 mol% of the total gas stream is condensed thereby forming a two-phase stream. The optimal mole

percentage will be dependant upon the composition ofthe gas undergoing liquefaction and other

process-related parameters readily ascertained by one possessing ordinary skilled in the art.

In one embodiment, the desired two-phase stream is obtained by cooling the entire

feed stream to such extent that the desired liquids percentage is obtained. In the preferred

embodiment, the gas stream is first cooled to near the liquefaction temperature and is then split

into a first stream and a second stream. The first stream undergoes additional cooling and partial

condensation and is then combined with the second stream thereby producing a two-phase stream

containing the desired percentage of liquids. This latter approach is preferred because ofthe

associated ease of operation and process control.

The two-phase stream is then fed to the upper section of a column wherein the

stream contacts the rising vapor stream from the lower portion of the column thereby producing

a heavies-rich liquid stream which functions as a reflux stream and a heavies-depleted vapor

stream which is produced from the column. As used herein, "heavies" will refer to any

predominantly organic compound possessing a molecular weight greater than ethane. The

column is unique in that it does not, as previously noted, employ a condenser for reflux

generation and further, does not employ a reboiler for vapor generation.

As previously noted, a methane-rich stripping gas stream is fed to the column.

This stream preferably originates from an upstream location where the methane-based gas stream

undergoing cooling has undergone some degree of cooling and liquids removal. Prior to

introduction into the base ofthe column, this gas stream is cooled via indirect contact, preferably

in a countercurrent manner, with the liquid product produced from the bottom ofthe column

thereby producing a warmed heavies-rich stream and a cooled methane-rich stripping gas stream. The methane-rich stripping gas may undergo partial condensation upon cooling and the resulting

cooled methane-rich stripping gas containing two phases may be fed directly to the column.

The employment ofthe cooled methane-rich stripping gas which contains small

amounts of C₃+ components in lieu of vapor generated from a reboiler which contains substantial

amounts of C₃+ components significantly reduces problems associated with fluids in the column

approaching critical conditions whereupon poor component separation results. This factor

becomes particularly significant when operating in the more preferred pressure range of about

550 to about 675 psia. The critical temperature and pressure of methane is -116.4°F and 673.3

psia. The critical temperature and pressure of propane is 206.2 °F and 617.4 psia and the critical

temperature and pressure of n-butane is 305.7°F and 551.25. The presence of appreciable

quantities of C₃+ components will (1) lower the critical pressure thereby approaching the

preferred operating pressures ofthe process and (2) raise the critical temperature. The resulting

effect is to make the separation ofthe components via vapor/liquid contacting more difficult. A

second factor distinguishing the uses ofthe cooled methane-rich stripping gas over vapor from a

reboiler is the temperature difference between these respective streams and the liquid effluent

from the last stage. Because it is preferred that the cooled methane-rich stripping gas be warmer

than the analogous vapor from a reboiler, this preferred stream possesses a greater ability to strip

the liquid phase ofthe lighter components. A temperature difference between the effluent liquid

from the column and the effluent stripping gas to the column is more preferably 20°F to 1 10°F,

still more preferably 40°F to 90°F, most preferably about 60°F to about 80°F.

The number of theoretical trays in the column will be dependant upon the

composition, temperature and flowrate ofthe inlet vapor stream to the column and the

composition, temperature, flowrate and liquid to vapor ratio ofthe two-phase stream fed to the upper section ofthe column. Such determination is readily within the abilities of one possessing

ordinary skill in the art. The theoretical number of trays may be provided via various types of

column packing (pall rings, saddles etc) or distinct contact stages (ex. trays) situated in the

column or a combination thereof. Generally, two (2) to fifteen (15) theoretical stages are

required, more preferably three (3) to ten (10), still more preferably four (4) to eight (8), and

most preferably about five (5) theoretical stages. Trays are generally preferred when the column

diameter is greater than six (6) ft.

Preferred Open-Cvcle Embodiment of Cascaded Liquefaction Process

The flow schematic and apparatus set forth in FIGURES 1 and 2 is a preferred

embodiment ofthe open-cycle cascaded liquefaction process and is set forth for illustrative

purposes. Purposely missing from the preferred embodiment is a nitrogen removal system,

because such system is dependant on the nitrogen content ofthe feed gas. However as noted in

the previous discussion of nitrogen removal technologies, methodologies applicable to this

preferred embodiment are readily available to those skilled in the art. Presented in FIGURES 3

and 4 in greater detail for illustrative purposes is the inventive cryogenic column and in

particular, the methodology for cooling and controlling the temperature ofthe stripping gas being

fed to the cryogenic column. Those skilled in the art will also recognized that FIGURES 1-4 are

schematics only and therefore, many items of equipment that would be needed in a commercial

plant for successful operation have been omitted for the sake of clarity. Such items might

include, for example, compressor controls, flow and level measurements and corresponding

controllers, additional temperature and pressure controls, pumps, motors, filters, additional heat exchangers, valves, etc. These items would be provided in accordance with standard engineering

practice.

To facilitate an understanding of FIGURES 1, 2, 3 and 4, items numbered 1 thru

99 generally correspond to process vessels and equipment directly associated with the

liquefaction process. Items numbered 100 thru 199 correspond to flow lines or conduits which

contain methane in major portion. Items numbered 200 thru 299 correspond to flow lines or

conduits which contain the refrigerant ethylene or optionally, ethane. Items numbered 300 thru

399 correspond to flow lines or conduits which contain the refrigerant propane. To the extent

possible, the numbering system employed in FIGURE 1 has been employed in FIGURES 2, 3,

and 4. In addition, the following numbering system has been added for additional elements not

illustrated in FIGURE 1. Items numbered 400 thru 499 correspond to additional flow lines or

conduits. Items numbered 500 thru 599 correspond to additional process equipment such as

vessels, columns, heat exchange means and valves, including process control valves. Items

numbered 600 thru 799 generally concern the process control system, exclusive of control

valves, and specifically includes sensors, transducers, controllers and setpoint inputs.

In almost all control systems, some combination of electrical, pneumatic or

hydraulic signals are used. However, the use of any other type of signal transmission compatible

with the process and equipment in use is withing the scope of this invention. With regard to the

invention depicted in FIGURES 1 through 4, lines designated as signal lines are depicted as dash

lines in the drawings. These lines are preferably electrical or pneumatic signal lines. Generally

the signals provided from any transducer are electric in form. However, the signals provided

from flow sensors are generally pneumatic in form. The transducing of these signals is not

always illustrated for the sake of simplicity because it is well known in the art that if a flow is measured in pneumatic form it must be transduced to electric form if it is to be transmitted in

electrical form by a flow transducer.

Referring to FIGURE 1 , gaseous propane is compressed in multistage compressor

18 driven by a gas turbine driver which is not illustrated. The three stages of compression

preferably exist in a single unit although each stage of compression may be a separate unit and

the units mechanically coupled to be driven by a single driver. Upon compression, the

compressed propane is passed through conduit 300 to cooler 20 where it is liquefied. A

representative pressure and temperature ofthe liquefied propane refrigerant prior to flashing is

about 100°F and about 190 psia. Although not illustrated in FIGURE 1, it is preferable that a

separation vessel be located downstream of cooler 20 and upstream of a pressure reduction

means, illustrated as expansion valve 12, for the removal of residual light components from the

liquefied propane. Such vessels may be comprised of a single-stage gas-liquid separator or may

be more sophisticated and comprised of an accumulator section, a condenser section and an

absorber section, the latter two of which may be continuously operated or periodically brought

on-line for removing residual light components from the propane. The stream from this vessel

or the stream from cooler 20, as the case may be, is passed through conduit 302 to a pressure

reduction means, illustrated as expansion valve 12, wherein the pressure ofthe liquefied propane

is reduced thereby evaporating or flashing a portion thereof. The resulting two-phase product

then flows through conduit 304 into high-stage propane chiller 2 wherein gaseous methane

refrigerant introduced via conduit 152, natural gas feed introduced via conduit 100 and gaseous

ethylene refrigerant introduced via conduit 202 are respectively cooled via indirect heat exchange

means 4, 6 and 8 thereby producing cooled gas streams respectively produced via conduits 154,

102 and 204. The gas in conduit 154 is fed to main methane economizer 74 which will be discussed in greater detail in a subsequent section and wherein the stream is cooled via indirect

heat exchange means 98. The resulting cooled compressed methane recycle stream produced via

conduit 158 is then combined with the heavies depleted vapor stream in conduit 120 from the

heavies removal column 60 and fed to the methane condenser 68.

The propane gas from chiller 2 is returned to compressor 18 through conduit 306.

This gas is fed to the high stage inlet port of compressor 18. The remaining liquid propane is

passed through conduit 308, the pressure further reduced by passage through a pressure reduction

means, illustrated as expansion valve 14 , whereupon an additional portion ofthe liquefied

propane is flashed. The resulting two-phase stream is then fed to chiller 22 through conduit 310

thereby providing a coolant for chiller 22. The cooled feed gas stream from chiller 2 flows via

conduit 102 to a knock-out vessel 10 wherein gas and liquid phases are separated. The liquid

phase which is rich in C₃₊ components is removed via conduit 103. The gaseous phase is

removed via conduit 104 and then split into two separate streams which are conveyed via

conduits 106 and 108. The stream in conduit 106 is fed to propane chiller 22. The stream in

conduit 108 becomes the feed to heat exchanger 62 and is ultimately the stripping gas to the

heavies removal column 60. Ethylene refrigerant from chiller 2 is introduced to chiller 22 via

conduit 204. In chiller 22, the feed gas stream, also referred to herein as a methane-rich stream,

and the ethylene refrigerant streams are respectively cooled via indirect heat transfer means 24

and 26 thereby producing cooled methane-rich and ethylene refrigerant streams via conduits 1 10

and 206. The thus evaporated portion ofthe propane refrigerant is separated and passed through

conduit 311 to the intermediate-stage inlet of compressor 18. Liquid propane refrigerant from

chiller 22 is removed via conduit 314, flashed across a pressure reduction means, illustrated as

expansion valve 16, and then fed to third stage chiller 28 via conduit 316. As illustrated in FIGURE 1 , the methane-rich stream flows from the intermediate-

stage propane chiller 22 to the low-stage propane chiller/condenser 28 via conduit 110. In this

chiller, the stream is cooled via indirect heat exchange means 30. In a like manner, the ethylene

refrigerant stream flows from the intermediate-stage propane chiller 22 to the low-stage propane

chiller/condenser 28 via conduit 206. In the latter, the ethylene refrigerant is totally condensed

or condensed in nearly its entirety via indirect heat exchange means 32. The vaporized propane

is removed from the low-stage propane chiller/condenser 28 and returned to the low-stage inlet at

the compressor 18 via conduit 320. Although FIGURE 1 illustrates cooling of streams provided

by conduits 110 and 206 to occur in the same vessel, the chilling of stream 110 and the cooling

and condensing of stream 206 may respectively take place in separate process vessels (ex., a

separate chiller and a separate condenser, respectively). In a similar manner, the preceding

cooling steps wherein multiple streams were cooled in a common vessel (ex., chiller) may be

conducted in separate vessels. The former arrangement is a preferred embodiment because ofthe

cost of multiple vessels and the requirement of less plant space.

As illustrated in FIGURE 1 , the methane-rich stream exiting the low-stage

propane chiller is introduced to the high-stage ethylene chiller 42 via conduit 112. Ethylene

refrigerant exits the low-stage propane chiller 28 via conduit 208 and is preferably fed to a

separation vessel 37 wherein light components are removed via conduit 209 and condensed

ethylene is removed via conduit 210. The separation vessel is analogous to the vessel earlier

discussed for the removal of light components from liquefied propane refrigerant and may be a

single-stage gas-liquid separator or may be a multiple stage operation which provides greater

selectivity in the removal of light components from the system. The ethylene refrigerant at this

location in the process is generally at a temperature of about -24 °F and a pressure of about 285 psia. The ethylene refrigerant via conduit 210 then flows to the ethylene economizer 34

wherein it is cooled via indirect heat exchange means 38 and removed via conduit 21 1 and

passed to a pressure reduction means illustrated as an expansion valve 40 whereupon the

refrigerant is flashed to a preselected temperature and pressure and fed to the high-stage ethylene

chiller 42 via conduit 212. Vapor is removed from this chiller via conduit 214 and routed to the

ethylene economizer 34 wherein the vapor functions as a coolant via indirect heat exchange means 46. The ethylene vapor is then removed from the ethylene economizer via conduit 216

and feed to the high-stage inlet on the ethylene compressor 48. The ethylene refrigerant which is

not vaporized in the high-stage ethylene chiller 42 is removed via conduit 218 and returned to

the ethylene economizer 34 for further cooling via indirect heat exchange means 50, removed

from the ethylene economizer via conduit 220 and flashed in a pressure reduction means

illustrated as expansion valve 52 whereupon the resulting two-phase product is introduced into

the low-stage ethylene chiller 54 via conduit 222.

Removed from high-stage ethylene chiller 42 via conduit 116 is a methane-rich

stream. This stream is then condensed in part via cooling provided by indirect heat exchange

means 56 in low-stage ethylene chiller 54 thereby producing a two-phase stream which flows via

conduit 118 to the benzene/aromatics/heavies removal column 60. As previously noted, the

methane-rich stream in conduit 104 was split so as to flow via conduits 106 and 108. The

contents of conduit 108 which is referred to herein as the methane-rich stripping gas is first fed to

heat exchanger 62 wherein this stream is cooled via indirect heat exchange means 66 thereby

becoming a cooled methane-rich stripping gas stream which then flows by conduit 109 to the

benzene heavies removal column 60. Liquid containing a significant concentration of benzene,

other aromatics and/or heavier hydrocarbon components is removed from the benzene/heavies removal column 60 via conduit 1 14, preferably flashed via a flow control means which can also

function as a pressure reduction means 97, preferably a control valve, and transported to heat

exchanger 62 by conduit 1 17. Preferably, the stream flashed via flow control means 97 is flashed

to a pressure about or greater than the pressure at the high stage inlet port to the methane

compressor. Flashing also imparts greater cooling capacity to said stream. In the heat exchanger

62, the stream delivered by conduit 117 provides cooling capabilities via indirect heat exchange

means 64 and exits said heat exchanger via conduit 119. In the benzene/aromatics/heavies

removal column 60, the two-phase stream introduced via conduit 118 is contacted with the

cooled methane-rich stripping gas stream introduced via conduit 109 in a countercurrent manner

thereby producing a benzene/heavies-depleted, methane-rich vapor stream via conduit 120 and a

benzene/heavies-enriched liquid stream via conduit 117.

The stream in conduit 119 is rich in benzene, other aromatics and/or other heavier

hydrocarbon components. This stream is subsequently separated into liquid and vapor portions

or preferably is flashed or fractionated in vessel 67. In each case a liquid stream rich in benzene,

other aromatics and or heavier hydrocarbon components and is produced via conduit 123 and a

second methane-rich vapor stream is produced via conduit 121. In the preferred embodiment

which is illustrated in FIGURE 1, the stream in conduit 121 is subsequently combined with a

second stream delivered via conduit 128 and the combined stream fed via conduit 140 to the high

pressure inlet port on the methane compressor 83. As previously noted, the gas in conduit 154 is fed to main methane economizer 74

wherein the stream is cooled via indirect heat exchange means 98. The resulting cooled

compressed methane recycle or refrigerant stream in conduit 158 is combined in the preferred

embodiment with the heavies depleted vapor stream from the heavies removal column 60 delivered via conduit 120 and fed to the low-stage ethylene condenser 68. In the low-stage

ethylene condenser, this stream is cooled and condensed via indirect heat exchange means 70

with the liquid effluent from the low-stage ethylene chiller 54 which is routed to the low-stage

ethylene condenser 68 via conduit 226. The condensed methane-rich product from the low-

stage condenser is produced via conduit 122. The vapor from the low-stage ethylene chiller 54

withdrawn via conduit 224 and low-stage ethylene condenser 68 withdrawn via conduit 228 are

combined and routed via conduit 230 to the ethylene economizer 34 wherein the vapors function

as coolant via indirect heat exchange means 58. The stream is then routed via conduit 232 from

the ethylene economizer 34 to the low-stage side ofthe ethylene compressor 48.

As noted in FIGURE 1, the compressor effluent from vapor introduced via the

low-stage side is removed via conduit 234, cooled via inter-stage cooler 71 and returned to

compressor 48 via conduit 236 for injection with the high-stage stream present in conduit 216.

Preferably, the two-stages are a single module although they may each be a separate module and

the modules mechanically coupled to a common driver. The compressed ethylene product from

the compressor is routed to a downstream cooler 72 via conduit 200. The product from the

cooler flows via conduit 202 and is introduced, as previously discussed, to the high-stage

propane chiller 2

The liquefied stream in conduit 122 is generally at a temperature of about -125°F

and a pressure of about 600 psi. This stream passes via conduit 122 through the main methane

economizer 74, wherein the stream is further cooled by indirect heat exchange means 76 as

hereinafter explained. From the main methane economizer 74 the liquefied gas passes through

conduit 124 and its pressure is reduced by a pressure reduction means which is illustrated as

expansion valve 78, which of course evaporates or flashes a portion ofthe gas stream. The flashed stream is then passed to methane high-stage flash drum 80 where it is separated into a gas

phase discharged through conduit 126 and a liquid phase discharged through conduit 130. The

gas-phase is then transferred to the main methane economizer via conduit 126 wherein the vapor

functions as a coolant via indirect heat transfer means 82. The vapor exits the main methane

economizer via conduit 128 where it is combined with the gas stream delivered by conduit 121.

These streams are then fed to the high pressure inlet port of compressor 83.

The liquid phase in conduit 130 is passed through a second methane economizer

87 wherein the liquid is further cooled by downstream flash vapors via indirect heat exchange

means 88. The cooled liquid exits the second methane economizer 87 via conduit 132 and is

expanded or flashed via pressure reduction means illustrated as expansion valve 91 to further

reduce the pressure and at the same time, vaporize a second portion thereof. This flash stream is

then passed to intermediate-stage methane flash drum 92 where the stream is separated into a gas

phase passing through conduit 136 and a liquid phase passing through conduit 134. The gas

phase flows through conduit 136 to the second methane economizer 87 wherein the vapor cools

the liquid introduced to 87 via conduit 130 via indirect heat exchanger means 89. Conduit 138

serves as a flow conduit between indirect heat exchange means 89 in the second methane

economizer 87 and the indirect heat transfer means 95 in the main methane economizer 74. This

vapor leaves the main methane economizer 74 via conduit 140 which is connected to the

intermediate stage inlet on the methane compressor 83.

The liquid phase exiting the intermediate stage flash drum 92 via conduit 134 is

further reduced in pressure by passage through a pressure reduction means illustrated as a

expansion valve 93. Again, a third portion ofthe liquefied gas is evaporated or flashed. The

fluids from the expansion valve 93 are passed to final or low stage flash drum 94. In flash drum 94, a vapor phase is separated and passed through conduit 144 to the second methane economizer

87 wherein the vapor functions as a coolant via indirect heat exchange means 90, exits the

second methane economizer via conduit 146 which is connected to the first methane economizer

74 wherein the vapor functions as a coolant via indirect heat exchange means 96 and ultimately

leaves the first methane economizer via conduit 148 which is connected to the low pressure port

on compressor 83.

The liquefied natural gas product from flash drum 94 which is at approximately

atmospheric pressure is passed through conduit 142 to the storage unit. The low pressure, low

temperature LNG boil-off vapor stream from the storage unit and optionally, the vapor returned

from the cooling ofthe rundown lines associated with the LNG loading system, is preferably

recovered by combining such stream or streams with the low pressure flash vapors present in

either conduits 144, 146, or 148; the selected conduit being based on a desire to match vapor

stream temperatures as closely as possible.

As shown in FIGURE 1, the high, intermediate and low stages of compressor 83

are preferably combined as single unit. However, each stage may exist as a separate unit where

the units are mechanically coupled together to be driven by a single driver. The compressed gas

from the low-stage section passes through an inter-stage cooler 85 and is combined with the

intermediate pressure gas in conduit 140 prior to the second-stage of compression. The

compressed gas from the intermediate stage of compressor 83 is passed through an inter-stage

cooler 84 and is combined with the high pressure gas in conduit 140 prior to the third-stage of

compression. The compressed gas is discharged from the high-stage methane compressor

through conduit 150, is cooled in cooler 86 and is routed to the high pressure propane chiller via

conduit 152 as previously discussed. FIGURE 1 depicts the expansion of the liquefied phase using expansion valves

with subsequent separation of gas and liquid portions in the chiller or condenser. While this

simplified scheme is workable and utilized in some cases, it is often more efficient and effective

to carry out partial evaporation and separation steps in separate equipment, for example, an

expansion valve and separate flash drum might be employed prior to the flow of either the

separated vapor or liquid to a propane chiller. In a like manner, certain process streams

undergoing expansion are ideal candidates for employment of a hydraulic expander as part ofthe

pressure reduction means thereby enabling the extraction of work energy and also lower

two-phase temperatures.

With regard to the compressor/driver units employed in the process, FIGURE 1

depicts individual compressor/driver units (i.e., a single compression train) for the propane,

ethylene and open-cycle methane compression stages. However in a preferred embodiment for

any cascaded process, process reliability can be improved significantly by employing a multiple

compression train comprising two or more compressor/driver combinations in parallel in lieu of

the depicted single compressor/driver units. In the event that a compressor/driver unit becomes

unavailable, the process can still be operated at a reduced capacity.

Preferred Embodiment of the Inventive Removal Process and Apparatus

Presented in FIGURE 2 is a preferred embodiment ofthe benzene, other aromatic

and/or heavier hydrocarbon component removal process and associated apparatus. As previously

noted, the two-phase stream fed to the benzene/aromatics/heavies removal column 60 via conduit

118 results from the cooling and partial condensing ofthe stream in conduit 116 via cooling

provided by heat exchange means 56 in ethylene chiller 54. In one embodiment, the entire stream in conduit 116 is cooled. In a preferred embodiment illustrated in FIGURE 2, the two-

phase stream is obtained by cooling and partially condensing a portion ofthe stream in conduit

116 and this portion is then combined with the remaining portion ofthe stream originating via

conduit 1 16.

Referring to FIGURE 2, the stream delivered via conduit 116 is split into a first

stream flowing in conduit 450 and a second stream flowing in conduit 452. The stream in

conduit 532 flows through an optional valve 532, preferably a hand control valve, to conduit 454

which delivers the first stream to ethylene chiller 54 wherein the stream undergoes at least partial

condensation via indirect heat exchange means 56 and exits said means via conduit 458. The

second stream in conduit 452 flows through a valve 530, preferably a control valve, into conduit

456 which is then combined with the first stream delivered via conduit 458. The combined

streams, now a two-phase stream, is delivered to column 60 via conduit 1 18. From an

operational perspective, the length of conduit 118 should be sufficient to insure adequate mixing

ofthe two streams such that equilibrium conditions are approached. The amount of liquids in the

two-phase stream in conduit 118 is preferably controlled via maintaining the streams at a desired

temperature. This is accomplished in the following manner. A temperature transducing device

688 in combination with a sensing device such as a thermocouple situated in conduit 118

provides an input signal 686 to a temperature controller 682. Also provided to the controller by

operator or computer algorithm is a setpoint temperature signal 684. The controller 682 responds

to the differences in the two inputs and transmits a signal 680 to the flow control valve 530

which is situated in a conduit wherein flows the portion ofthe stream delivered via conduit 116

which does not undergo cooling via heat exchanger means 56 in chiller 54. The transmitted signal 680 is scaled to be representative ofthe position ofthe control valve 530 required to

maintain the flowrate necessary to obtain the desired temperature in conduit 118.

These feedstreams to the process step wherein benzene, other aromatic and/or

heavy hydrocarbon components are removed are the two-phase process stream from ethylene

chiller 54 delivered via conduit 118 to the upper section of column 60 and the methane-rich

stripper gas delivered via conduit 108. Although depicted in FIGURE 1 as originating from the

feed gas stream from the first stage of propane cooling, this stream can originate from any

location within the process or may be an outside methane-rich stream. As illustrated in FIGURE

2, at least a portion of the methane-rich stripper gas undergoes cooling in heat exchanger 62 via

indirect heat exchange means 62 prior to entering the base of column 60. Effluent streams from this inventive process step are the heavies-depleted gas stream from column 60 produced via

conduit 120 and the warmed heavies-rich stream produced via conduit 119. As illustrated in

FIGURE 2, a heavy-rich stream is produced from column 60 and undergoes warming in heat

exchanger 62 via indirect heat exchange means 66. It is in this manner that the column effluent

produced via conduit 114 cools the stripping gas fed to the column via conduit 109.

The number of theoretical stages in column 60 will be dependent on the

composition ofthe feedstreams to the column. Generally, two (2) to fifteen (15) theoretical

stages will be required. The preferred number of stages is three (3) to ten (10), still more

preferably is four (4) to eight (8) and from an operational and cost perspective, the most

preferred number is about five (5). The theoretical stages may be made available via packing,

plates/trays or a combination thereof. Generally, packing is preferred in columns of less than

about six (6) ft. diameter and plates/trays on columns of greater than about six (6) ft. diameter.

As illustrated in FIGURE 2, the upper section of column wherein the two-phase stream in conduit 1 18 is fed is designed to facilitate gas/liquid separation. The top ofthe column

preferably contains a means for demisting or removing entrained liquids from the vapor stream.

This means is to be located between the point of entry of conduit 118 and the point of exit of

conduit 120.

As illustrated in FIGURE 2, the heavies-rich liquid stream produced via conduit

114 flows through control valve 97 and conduit 117 to heat exchanger 62 wherein said stream

provides cooling via indirect heat transfer means 64 and is produced from heat exchanger 62 via

conduit 119 as a warmed heavies-rich stream. Depending on the operational pressure of

downstream processes, the cooling ability of this stream can be enhanced by flashing to a lower

pressure upon flow through control valve 97. This process stream produced via conduit 119 may

be utilized directly or undergo subsequent treatment for the removal of lighter components. In

the preferred embodiment illustrated in FIGURE 2, the stream is fed to a demethanizer 67.

The flowrate of heavies-rich liquid from column 60 may be controlled via various

methodologies readily available to one skilled in the art. The control apparatus illustrated in

FIGURE 2 is a preferred apparatus and is comprised ofa level controller device 600, also a

sensing device, and a signal transducer connected to said level controller device, operably located

in the lower section of column 60. The controller 600 establishes an output signal 602 that either

typifies the flowrate in conduit 114 required to maintain a desired level in column 60 or indicates

that the actual level has exceeded a predetermined level. A flow measurement device and

transducer 604 operably located in conduit 114 establishes an output signal 606 that typifies the

actual flowrate ofthe fluid in conduit 114. The flow measurement device is preferably located

upstream ofthe control valve so as to avoid sensing a two-phase stream. Signal 602 is provided

as a set point signal to flow controller 608. Signals 602 and 608 are respectively compared in flow controller 608 and controller 608 establishes an output signal 614 responsive to the

difference between signals 602 and 606. Signal 614 is provided to control valve 97 and valve 97

is manipulated responsive to signal 614. A setpoint signal (not illustrated) representative of a

desired level in column 60 may be manually inputted to level controller 600 by an operator or in

the alternative, be under computer control via a control algorithm. Depending on the operating

conditions, operator or computing machine logic is employed to determine whether control will

be based on liquid level or flowrate. In response to the variable flowrate input of signal 606 and

the selected setpoint signal, the controller 608 provides an output signal 614 which is responsive

to the difference between the respective input and setpoint signals. This signal is scaled so as to

be representative, as the case may be, ofthe position ofthe control valve 97 required to maintain

the flowrate of fluid substantially equal to the desired flowrate or the liquid level substantially

equal to the desired liquid level, as the case may be.

In the heat exchanger 62, the heavies-rich stream, which cools the methane-rich

stripping gas stream, is routed to the heat exchanger via conduit 117. The heavies-rich stream

flows thru indirect heat exchange means 66 and is produced from the heat exchanger via conduit

119. The degree to which the methane-rich stripping gas is cooled by the heavies-bearing stream

prior to entry into the column may be controlled via various methodologies readily available to

one skilled in the art. In one embodiment, the entire methane-rich stripping gas stream is fed to

the heat exchanger and the degree of cooling controlled by such parameters as the amount of

heavies-rich liquid stream made available for heat transfer, the heat transfer surface areas

available for heat transfer and/or the residence times ofthe fluids undergoing heating or cooling

as the case may be. In a preferred embodiment, the methane-rich stripping gas stream delivered

via conduit 108 flows through control valve 500 into conduit 400 whereupon the stream is split and transferred via conduits 402 and 403. The stream flowing through conduit 403 ultimately

flows through indirect heat transfer means 64 in heat exchanger 62. A means for manipulating

the relative flowrates of fluid in conduits 402 and 403 is provided in either conduits 402 or 403

or both. The means illustrated in FIGURE 2 are simple hand control valves, designated 502 and

504, which are respectively attached to conduits 404 and 407. However, a control valve whose

position is manipulated by a controller and for which input to the controller is comprised ofa

setpoint and signal representative of flow in the conduit, such as that discussed above for the

heavies-bearing stream, may be substituted for one or both ofthe hand control valves. In any

event, the valves are operated such that the temperature approach difference ofthe streams in

conduits 117 and 404 to heat exchanger 62 does not exceed 50 °F whereupon damage to the heat

exchanger might result. The cooled fluid leaves the indirect heat transfer means 64 via conduit

405 and is combined at a junction point with uncooled methane-rich stripping gas delivered via

conduit 407 thereby forming the cooled methane-rich stripping gas stream which is delivered to

the column via conduit 109.

Operably located in conduit 109 is a flow transducing device 616 which in

combination with a flow sensing device such as an orifice plate (not illustrated) establishes an

output signal 618 that typifies the actual flowrate ofthe fluid in the conduit. Signal 618 is

provided as a process variable input to a flow controller 620. Also provided either manually or

via computer output is a set point value for the flowrate represented by signal 622. The flow

controller then provides an output signal 624 which is responsive to the difference between the

respective input and setpoint signals and which is scaled to be representative ofthe position of

the control valve required to maintain the desired flowrate in conduit 109. In another embodiment, the relative flowrate of fluid through conduits 402 and

403 can be controlled via locating a temperature sensing device and a transducer connected to

said device, if so required, in conduit 109 and using the resulting output and a setpoint

temperature as input to a flow controller which would generate an output signal responsive to the

difference in the two signals and scaled to be representative of a control valve position required

to maintain the desired flowrate in conduit 109. Such control valves could be substituted for

hand valves 502 and/or 504.

In still yet another embodiment depicted in FIGURE 3, the temperature ofthe

stripping gas to column 60 is controlled in the following manner. Temperature transducer 704 in

combination with a measuring device such as a thermocouple operably located in conduit 117

provides an output signal 708 which is representative ofthe actual temperature of liquid flowing

in conduit 117. Signal 708 is provided as a first input to the ratio calculator 700. Ratio

calculator 700 is also provided with a second temperature signal 706 representative ofthe

temperature of fluid flowing into conduit 109. Signal 706 originates in temperature transducer

702 whose output signal 706 is responsive to a sensing element such as a thermocouple operably

located in conduit 109. In response to signals 706 and 708 ratio calculator 700 provides an

output signal 710 which is representative ofthe ratio of signals 706 and 708. Signal 710 is

provided as an input to ratio controller 712. Ratio controller 712 is also provided with a set point

signal 714 which is representative ofthe desired temperature ratio for the fluids flowing in

conduits 109 and 114. Responsive to signals 710 and 714, ratio controller 712 provides an

output signal 716 which is responsive to the difference between signals 710 and 714. Signal 716

is scaled to be representative ofthe position of control valve 534, which is operably located in by-pass conduit 718, required to maintain the desired ratio represented by set point signal 714.

Control valve 534 is manipulated responsive to signal 716.

In accordance with the most preferred control methodology depicted in FIGURE

4 where like reference numerals are used for elements shown in the previous Figures, an

automatic start-up of column 60 is facilitated by high selector 728. It is noted that the set point

724 of temperature controller 722 is desirably set at a temperature compatible with the liquid in

the column 60. On start-up however, the temperature in conduit 109 will be at or near ambient

temperature. Accordingly connecting signal 726 directly to manipulate valve 536 would cause

valve 536 to close and not allow flow ofthe warm dry gas to a cryogenic separation column 60

during startup. This problem is overcome by temporarily selecting signal 742 to manipulate

valve 536 as described below.

Responsive to signals 706 and 724 temperature controller 722 provides an output

signal 726 responsive to the difference between signals 706 and 724. Signal 726 is scaled to be

representative ofthe position of control valve 536 which is operably located in conduit 108

required to maintain the actual temperature ofthe fluid in conduit 109 substantially equal to the

desired temperature representative by signal 724. As previously stated, however, the desired

value for set point signal 724 will not allow start-up ofthe column. Accordingly signal 726 is

provided to a signal selector 728. Signal selector 728 is also provided with a control signal 742

which is responsive to the difference between signals 736 and 740 and is scaled to be

representative of the position of control valve 536 required to maintain the temperature of fluid

in conduit 119 substantially equal to the desired temperature represented by signal 740. On start¬

up ofthe column, the actual temperature of fluid in conduit 119 will be less than the desired

temperature represented by signal 740. Accordingly, connecting signal 742 to valve 536 would cause valve 536 to open so as to lower the temperature represented by signal 706. High selector

728 decides which ofthe control signals 726 and 742 manipulate the valve 536.

Start-up proceeds like this. Feed gas is introduced into the top ofthe cryogenic

separation column 60 in the upper section. When the temperature ofthe feed gas cools to the

condensing temperature ofthe impurity to be removed, liquid begins to build a level in the column 60. Level controller 600 senses the level and its output opens valve 97 responsive to

signal 614. Low temperature liquid is then passed to heat exchanger 62 and exchanges heat with

a warm dry gas stream through conduit 108 and valve 536. Valve 536 is initially opened by

signal 742 on set point temperature. After dry gas flow is initiated temperature transducer 702

senses a sharply colder temperature resulting in signal 726 being selected by the high selector

728. The start-up controls assist the operator in providing a smooth safe start-up and reduce the

level of human attention required.

The warmed heavies-rich liquid stream from heat exchanger 62 is fed via conduit

119 to the demethanizer column 67 which contains both rectifying and stripping sections. The

rectifying and stripping sections may contain distinct stages (e.g., trays, plates) or may provide

for continuous mass transfer via column packing (eg., saddles, racking rings, woven wire) or a

combination ofthe preceding. Generally, packing is preferred for columns possessing a diameter

of less than about six (6) ft and distinct stages preferred for columns possessing a diameter of

greater than about six (6) ft. The number of theoretical stages in both the rectifying and stripping

sections is dependant on the desired composition of the final products and the composition of the

feed stream. Preferably the stripping or lower section contains 4 to 20 theoretical stages, more

preferably 8 to 12 theoretical stages, and most preferably about 10 theoretical stages. In a similar

manner, the upper or rectifying section ofthe column preferably contains 4 to 20 theoretical stages, more preferably 8 to 13 theoretical stages, and most preferably about 10 theoretical

stages.

A conventional reboiler 524 is provided at the bottom to provide stripping vapor.

In the preferred embodiment presented in FIGURE 2, liquid from the lower-most stage in the

demethanizer is provided to the reboiler via conduit 428 wherein said fluid is heated via an

indirect heat transfer means 525 with a heating medium delivered via conduit 440 and returned

via conduit 442 which is connected to flow control valve 526 which is in turn connected to

conduit 444. Vapor from the reboiler is returned to the demethanizer column via conduit 430

and liquids are removed from the reboiler via conduit 432. Said stream in conduit 432 may

optionally be combined in conduit 436 with a second liquids stream produced from the bottom of

the demethanizer via optional conduit 434. The total liquids stream produced from the

demethanizer via conduits 436 and/or 432, as the case may be, may optionally flow thru cooler

520 and produced via conduit 438. A means for controlling liquid flow is inserted into one or

both ofthe preceding conduits. In one embodiment as illustrated in FIGURE 2, the flow control

means is comprised of control valve 522 which is inserted between conduits 438 and 123. The

position ofthe control valve 522 is manipulated by a flow controller 632 which is responsive to

the differences between a setpoint input signal 628 from a level control device 626 and the actual

flowrate of fluid in conduit 438 represented by signal 631. A set point flowrate 630 for level

controller 626 may be provided via operator or computer algorithm input. Output from the

controller 632 is signal 634 which is scaled to be representative ofthe position ofthe control

valve 522 required to maintain the desired flowrate in conduit 438 to maintain the desired level in 67. Although various control techniques are readily available for regulating the

flowrate of stripping vapor to the column 67 via conduit 430, a preferred technique is based on

the temperature ofthe return vapor. A temperature transducing device 636 in combination with a

sensing device such as a thermocouple situated in conduit 430 provides an input signal 638 to a

temperature controller 642. Also provided to the controller by operator or computer algorithm is

a setpoint temperature signal 640. The controller 642 responds to the differences in the two

inputs and transmits a signal 644 to the flow control valve 526 which is situated in a conduit

containing the heating medium, preferably conduits 440 or 444, most preferably conduit 444 as

illustrated. The transmitted signal 644 is scaled to be representative ofthe position ofthe control valve 526 required to maintain the flowrate necessary to obtain the desired temperature in

conduit 440.

A novel aspect ofthe demethanizer column is the manner in which reflux liquids

are generated. As illustrated in FIGURE 2, the overhead product exits the demethanizer column

67 via conduit 410 whereupon at least a portion of said stream is partially condensed upon

flowing through indirect heat exchange means 510 in heat exchanger 62 which is cooled via the

heavies-rich liquid product from the heavies removal column 60. In a preferred embodiment, the

heavies-rich liquid product is first employed for cooling of at least a portion ofthe overhead

vapor stream and then employed for cooling ofthe methane-rich stripping gas stream. The

condensed liquids resulting from cooling via the heavies-rich liquid stream become the source of

the reflux for demethanizer column 67. Preferably, the heat exchange between the two

designated streams occurs in a countercurrent manner. In one embodiment, the entire stream

may flow to heat exchanger 62 in the manner previously discussed for the cooling ofthe entire

methane stripping gas. In a preferred embodiment which is illustrated in FIGURE 2, the overhead vapor product in conduit 410 is split into streams flowing in conduits 412 and 414.

The stream in conduit 414 is cooled in heat exchanger 62 by flowing said stream through indirect

heat exchange means 510 in exchanger 62 and the resulting cooled stream is produced via

conduit 418. The relative flowrates ofthe vapor streams in conduits 412 and 414 or 418 are

controlled by a flow control means, preferably a flow control valve through which overhead

vapor may flow without flowing through the heat exchanger thereby avoiding the control ofa

two-phase fluid. Vapor flowing in conduit 412 flows through flow control means 512 and is

produced therefrom via conduit 416. Conduits 416 and 418 are then joined thereby resulting in a combined cooled two-phase stream which flows through conduit 420. Situated in conduit 420 is

a temperature transducing device 646, in combination with a temperature sensing device,

preferably a thermocouple, provides a signal 648 representative ofthe actual temperature ofthe

fluid flowing in conduit 420 to temperature controller 652. A desired temperature 650 is also

inputted to the controller 652 either manually or via a computational algorithm. Based on a

comparison ofthe input via the transducing device 646 and the setpoint 650, the controller 652

then provides an output signal 654 to the valve 512 which is scaled to manipulate the valve 512

in an appropriate manner such that the setpoint temperature is approached or maintained. The

resulting two-phase fluid in conduit 420 is then fed to separator 514 from which is produced a

methane-rich vapor stream via conduit 422 and a reflux liquid stream via conduit 424. In another

preferred embodiment, the preceding methodology is employed but the heavies-rich stream in

conduit 1 17 is first employed for cooling ofthe stream delivered via conduit 414 prior to cooling

the stream delivered via conduit 414. As illustrated in FIGURE 1, the methane rich vapor stream

in conduit 121 can be returned to the open methane cycle for subsequent liquefaction. The

pressure ofthe demethanizer and associated equipment is controlled by automatically manipulating control valve 518 responsive to a pressure transducer device 656 operably located

in conduit 422. The control valve is connected on the inlet side to conduit 422 and on the outlet

side to conduit 121 which preferably is directly or indirectly connected to the low pressure inlet

port on the methane compressor, the pressure transducing device 656 in combination with a

sensing device, provides a signal 658 to a pressure controller 660 which is representative ofthe

actual pressure in conduit 422. A set point pressure signal 662 is also provided as input to the

pressure controller 660. The controller then generates a response signal 664 representative of the

difference between the pressure sensing device signal 658 and the setpoint signal 662. Signal

664 is scaled in such a manner as to activate the valve 518 according for approach and

maintenance ofthe setpoint pressure. In one embodiment, the controller and control valve and

optionally, the pressure sensing transducer 656 are embodied in a single device commonly called

a back pressure regulator.

The reflux from the separator ultimately flows to the demethanizer. In the

preferred embodiment illustrated in FIGURE 2, the reflux leaves the separator 514 via conduit

424, flows thru pump 516, and then flows thru conduit 425 , control valve 519, and conduit 426

whereupon the stream is introduced into the upper section ofthe demethanizer column. In this

embodiment, the flowrate of reflux is controlled via input from a level control device 666 which

is responsive to a sensing device located in the lower section ofthe separator 514. Controller

666 generates a signal 668 representative ofthe flowrate in conduit 426 required to maintain the

desired level in separator 514, signal 668 is provided as a setpoint input to flow controller 670 to

which is also fed a signal 671 which typifies the actual flowrate in conduit 425. The controller

670 then generates a signal 674 to control valve 519 which is representative ofthe difference in signals and scaled to provide for appropriate liquids flow through the flow control valve 519

such that liquid level in separator 514 is controlled.

The controllers previously discussed may use the various well-known modes of

control such as proportional, proportional-integral, or proportional-integral-derivative (PID). A

digital computer having backup accommodations is preferred in the preferred embodiment

depicted in FIGURE 4 for calculating the required control signals based on measured process

variables as well as set points supplied to the computer. Any digital computer having software

that allows operation ofa real time environment for reading values of external variables and

transmitting signals to external devices is suitable for use. The PID controllers shown in

FIGURES 2, 3, and 4 can utilize the various modes of control such as proportional, proportional-

integral or proportional-integral-derivative. In the preferred embodiment a proportional-integral

mode is utilized. However, any controller having capacity to accept two or more input signals

and produce a scaled output signal representative ofthe comparison ofthe two input signals is

within the scope ofthe invention.

The scaling of an output signal by a controller is well known in the control

systems art. Essentially, the output ofa controller can be scaled to represent any desired factor

or variable. An example of this is where a desired temperature and an actual temperature are

compared by controller. The controller output might be a signal representative of a flow rate ofa

"control" gas necessary to make the desired and actual temperatures equal. On the other hand,

the same output signal could be scaled to represent a pressure required to make the desired and

actual temperatures equal. If the controller output can range from 0-10 units, then the controller

output signal could be scaled so that an output having a level of 5 units corresponds to 50%

percent or a specified flow rate or a specified temperature. The transducing means used to measure parameters which characterize a process in the various signals generated thereby may

take a variety of forms or formats. For example the control elements of this system can be

implemented using electrical analog, digital electronic, pneumatic, hydraulic, mechanical, or

other similar types of equipment or combination of such types of equipment.

Selective control loops are used in a variety of process situations for selecting an

appropriate control action. Typically a normal control signal is overridden by a secondary

control signal that has a higher priority in the event of certain process conditions. For example,

hazardous conditions can be avoided, or desirable features such as automatic start-up can be

implemented by temporarily selecting a secondary control signal.

The specific hardware and/or software utilized in such feedback control systems is well known in the field of process plant control. See for example Chemical Engineering's

Handbook, Sth Ed., McGraw-Hill, pgs. 22-1 to 22-147.

While specific cryogenic methods, materials, items of equipment and control

instruments are referred to herein, it is to be understood that such specific recitals are not to be

considered limiting but are included by way of illustration and to set forth the best mode in

accordance with the present invention.

EXAMPLE I

This Example shows via computer simulation the efficiency ofthe process

described in the specification for the removal of benzene and heavier components from a

methane-based stream immediately prior to liquefaction ofthe methane-based stream in major

portion. The flowrates are representative to those existing in a 2.5 million metric tonne/year

LNG plant employing the liquefaction technology set forth in FIGURES 1 and 2. The benzene

concentrations in the methane-based gas streams employed in this Example are considered to be representative of those possessed by many candidate natural gas streams at this location in the

process. However, the methane-based gas streams are considered to be relatively lean in the

heavier hydrocarbon components (i.e., C₃+). Simulation results were obtained using Hyprotech's

Process Simulation HYSIM, version 386/C2.10, Prop. Pkg PR/LK.

Presented in Table 1 are the compositions, temperatures, pressures and phase

conditions ofthe influent and effluent streams to the heavies removal column. The simulation is

based upon the column containing 5 theoretical stages. The partially condensed stream, also

referred to as the two-phase stream, which will latter undergo liquefaction in major proportion is

first fed to the uppermost stage in the column (Stage 1). The temperature of this stream is -

112.5 °F and the pressure is 587.0 psia. As previous noted, this stream has undergone partial

condensation such that the stream is 98.24 mol% vapor.

The cooled methane-rich stripping gas fed into the lowermost stage (Stage 5) is

taken from the upstream location depicted in FIGURE 1. This stream is cooled from

approximately 63 °F to -10°F via countercurrent heat exchange with the heavies-rich liquid

stream produced from Stage 5. During such heat exchange as depicted in FIGURE 2, this stream

is heated from approximately -78 °F to approximately 62 °F. This stream may also be employed

to cool the overhead vapors from the demethanizer column. Presented in Table 2 are the

simulated temperatures, pressures, and relative flowrates of each phase on a stagewise basis

within the column. Presented in Table 3 for each stage are the respective liquid and vapor

equilibrium compositions.

The warmed heavies-rich stream is then fed to the demethanizer column which

contains rectifying and stripping sections wherefrom is produced a methane/ethane rich stream which preferably is recycled back as feed to the high stage inlet port on the methane compressor

and a stream rich in natural gas liquids.

The efficiency ofthe process for aromatics/heavy removal is illustrated by a

comparison ofthe combined nitrogen, methane and ethane mole percentages in the feed streams

to Stages 1 and 5 and the product from Stage 1. These percentages for each stream are

respectively 99.88, 99.89 and 99.94 mol percent. The process therefore produces a product

stream richer in these light components than either ofthe two gaseous feed streams.

The efficiency of the process for benzene and heavier aromatics removal is

illustrated by a comparison ofthe enrichment ratios which is defined to be the mole percent of

said component in the liquid product from Stage 5 divided by the mole percent of said

component in the vapor product from Stage 1. Using benzene as an example, the respective mole

fractions are 0.1616E-04 and 0.00352. This results in an enrichment ratio of approximately 220.

An additional basis for illustrating the efficiency of the process are the enrichment

ratios for the C3+ components in the feed streams to Stages 1 and 5 and the liquid product stream

produced from Stage 1. This ratio varies from about 45 for propane to about 200 for n-octane.

The respective ratios between the product streams varies from about 50 for propane to about

20,000 for n-octane.

EXAMPLE μ This Example, like that previously presented, shows via computer simulation the

efficiency ofthe process described in the specification for the removal of benzene and heavier

components from a methane-based gas stream immediately prior to liquefaction ofthe stream in

major portion. The flowrates are representative of those existing in a 2.5 million metric

tonne/year LNG plant employing the liquefaction technology set forth in FIGURES 1 and 2. The benzene concentrations in the methane-rich feed streams employed in this Example are

considered to be representative ofthe concentrations existing for many candidate gas streams at

this location in the process. However, the concentrations of ethane and heavier components in

the gas stream have been increased significantly thereby representing a richer gas stream and

placing a greater burden on the process for the removal of both these components and benzene.

This example illustrates in greater detail the ability ofthe process to simultaneously remove

benzene and heavier hydrocarbon components. In addition, this Example illustrates the ability of

the benzene removal process to tolerate significant process upsets in the form of significant

increases in ethane and heavier hydrocarbon concentrations without significantly affecting the

efficiency and operability ofthe benzene removal process. Furthermore, this example illustrates the ability ofthe process to recover heavies hydrocarbons as a separate liquefied stream.

Simulation results were obtained using Hyprotech's Process Simulation HYSIM, version

386/C2.10, Prop. Pkg PR/LK.

Presented in Table 4 are the compositions, temperatures, pressures and phase conditions ofthe influent and effluent streams to the heavies removal column. The simulation is

based upon the column containing 5 theoretical stages. The partially condensed stream, also

referred to as the two-phase stream, which will undergo liquefaction in major proportion is first

fed to the uppermost stage in the column (Stage 1). The temperature of this stream is -91.2°F

and the pressure is 596.0 psia. As noted in the Specification, this stream has undergone partial

condensation such that the stream is 94.04 mol% vapor.

The methane-rich stripping stream fed into the lowermost stage (Stage 5) is taken

from the upstream location depicted in FIGURE 1. This stream is cooled from approximately - 10 F via countercurrent heat exchange with the liquid product stream produced from Stage 5. As

noted in Table 4, this stream has undergone partial condensation in the course of cooling.

Presented in Table 5 are the simulated temperatures, pressures, and relative

flowrates of each phase on a stagewise basis within the column. Presented in Table 6 for each

stage are the respective liquid and vapor equilibrium compositions.

The efficiency ofthe process for heavies removal is illustrated by a comparison of

the combined nitrogen, methane and ethane mole percentages in the feed streams respectively to

Stages 1 and 5 and the product stage from Stage 1. These percentages are respectively 97.85,

97.30, and 99.37 mol percent. The process produces a product stream significantly richer in

these components than either ofthe two gaseous feed streams.

The efficiency ofthe process for benzene and heavier aromatics removal is

illustrated by a comparison ofthe enrichment ratios which for benzene is as defined in Example

1. The respective mole fractions are 0.003E-04 and 0.00923 thus resulting in an enrichment

ratio of approximately 30. An additional basis for illustrating the efficiency ofthe process are the enrichment

ratios for the C3+ components in the feed streams to Stages 1 and 5 and the liquid product stream

produced from Stage 1. This ratio varies from about 19 for propane to about 30 for n-octane.

The respective ratios between the product streams varies from about 67 for propane to about

19,000 for n-octane. TABLE 1

FEEDSTREAM AND SIMULATED PRODUCT STREAM COMPOSITIONS AND PROPERTIES

Feed Streams¹ Product Streams'

Stage 1 Stage 5 Stage 1 Stage 5

Nitrogen 0.0022 0.0007 0.002169 0.000107

C0₂ 0.7587E-04 0.8806E-04 0.000075 0.000279

Methane 0.9726 0.9686 0.974167 0.559178

Ethane 0.0242 0.0296 0.023043 0.357346

Ethylene 0.0000 0.0000 0.000000 0.000000

Propane 0.0005 0.0006 0.000404 0.026993 i-Butane 0.8998E-04 0.0001 0.000055 0.009050 n-Butaπe 0.0001 0.0001 0.000059 0.013291 i-Pentane 0.3442E-04 0.4031E-04 0.00001 1 0.006026 n-Pentane 0.3340E-04 0.4031E-04 0.881E-05 0.006391 n-Hexane 0.2424E-04 0.3023E-04 0.257E-05 0.005627 n-Heptane 0.3230E-04 0.4031E-04 0.125E-05 0.008054 n-Octane 0.1615E-04 0.2015E-04 0.221E-06 0.004132

Benzene 0.1616E-04 0.2015E-04 0.258E-05 0.003526 n-Nonane 0.0000 0.0000 0.000000 0.000000

Temperature -112.45°F -10.00°F -1 12.32°F -78.09°F

Pressure 587.01 psia 601.00 psia 587.00 psia 589.00 psia

Vapor % 98.24% 100% 100% 0.00%

Flowrate (lb mole/hr) 60347.00 1203.0 6131 1.53 238.46

'Compositions are on mole fraction basis. TABLE 2

SIMULATION RESULTS OF FLOW CHARACTERISTICS AND FLUID PROPERTIES WITHIN THE COLUMN

Flow Rates (lb mole/hr)

Stage No. Pressure Temperature psia °F Product

Liquid Vapor Feed Streams

1 587.0 -1 12.3 1060.3 60347.0' 6131 1.5²

2 587.5 -108.2 917.8 2024.9

3 588.0 -101.1 761.5 1882.4

4 588.5 -90.8 619.0 1726.1

5 589.0 -78.1 1583.5 1203.0³ 238.5⁴

'Feed to Stage 1 is 98.24 mol % vapor.

²Product removed from Stage 1, 100 mol % vapor.

³Feed to Stage 5, 100 mol % vapor.

⁴Product removed from Stage 5, 0 mol % vapor.

TABLE 3

SIMULATED LIQUID/VAPOR STREAM COMPOSITIONS LEAVING EACH THEORETICAL STAGE (Mole Fraction)

Nitrogen co₂ Methane Ethane Propane i-Butane n-Butane

Stage 1

Vapor 0.002169 0.00075 0.974167 0.023043 0.000404 0.000055 0.000055

Liquid 0.000772 0.000173 0.874962 0.105444 0.006229 0.002030 0.002965

Stage 2

Vapor 0.000811 0.000110 0.967766 0.030734 0.000436 0.000057 0.000059

Liquid 0.000263 0.000252 0.832784 0.145068 0.007288 0.002348 0.003425

Stage 3

Vapor 0.000565 0.000144 0.954226 0.044398 0.000514 0.000063 0.000064

Liquid 0.000159 0.000317 0.761049 0.211924 0.009202 0.002861 0.004152

Stage 4

Vapor 0.000547 0.000163 0.933571 0.064781 0.000745 0.000082 0.000080

Liquid 0.000131 0.000329 0.669188 0.295174 0.013204 0.003786 0.005372

Stage 5

Vapor 0.000571 0.000154 0.913194 0.084077 0.001548 0.000194 0.000191

Liquid 0.000107 0.000279 0.559178 0.357346 0.026933 0.009050 0.013291

TABLE 3

SIMULATED LIQUID/VAPOR STREAM COMPOSITIONS

LEAVING EACH THEORETICAL STAGE (Mole Fraction)

(CONTINUED) i-Pentane n-Pentane n-Hexane n-Heptane n-Octane Benzene

Stage 1

Vapor 0.000011 8.81E-06 2.57E-06 1.25E-06 2.21E-07 2.58E-06

Liquid 0.001331 0.001408 0.001236 0.001768 0.000907 0.000775

Stage 2

Vapor 0.000011 8.54E-06 2.39E-06 1.12E-06 1.90E-07 2.35E-06

Liquid 0.001536 0.001625 0.001427 0.002042 0.001047 0.000894

Stage 3

Vapor 0.000011 8.64E-06 2.30E-06 1.03E-06 1.68E-07 2.17E-06

Liquid 0.001854 0.001961 0.001720 0.002461 0.01262 0.001078

Stage 4

Vapor 0.000014 0.000010 2.60E-06 1.14E-06 1.80E-07 2.31E-06

Liquid 0.002328 0.002446 0.002125 0.003031 0.001554 0.001332

Stage 5

Vapor 0.000033 0.000024 6.08E-06 2.57E-06 3.93E-07 4.83E-06

Liquid 0.006026 0.006391 0.005627 0.008054 0.004132 0.003526

TABLE 4

FEEDSTREAM AND SIMULATED PRODUCT STEAM COMPOSITIONS AND PROPERTIES (Mole Fraction)

Feed Streams¹ Product Streams¹

Stage 1 Stage 5 Stage 1 Stage 5

Nitrogen 0.0024 0.0006 0.002301 0.000060 co₂ 0.7074E-04 0.8851E-04 0.000072 0.000106

Methane 0.9478 0.9361 0.966005 0.346889

Ethane 0.0283 0.0363 0.025421 0.145714

Ethylene 0.0000 0.0000 0.000000 0.000000

Propane 0.0120 0.0145 0.005277 0.227598 i-Butane 0.0024 0.0030 0.000467 0.062744 n-Butane 0.0028 0.0036 0.000367 0.078635 i-Pentane 0.0010 0.0013 0.000049 0.030295 n-Pentane 0.0008 0.0011 0.000026 0.024383 n-Hexane 0.0013 0.0018 0.000012 0.043792 n-Heptane 0.0007 0.0010 0.170E-05 0.024376 n-Octane 0.0002 0.0003 0.1 1 1E-06 0.006019

Benzene 0.0003 0.0004 0.283E-05 0.009229 n-Nonane 0.4853E-05 0.6724E-05 0.851E-09 0.000160

Temperature -91.20°F -10.00°F -88.19°F -31.98°F

Pressure 596.01 psia 610 psia 596.00 psia 598.00 psia

Vapor % 94.04% 98.94% 100% 0.00%

Flowrate (lb mole/hr) 57109.78 7668.00 62724.19 2053.60

'Compositions are on mole fraction basis TABLE 5

SIMULATION RESULTS OF FLOW CHARACTERISTICS AND FLUID PROPERTIES WITHIN THE COLUMN

Fiow Rates (lb mole/hr)

Stage Pressure Temperature

No. psia °F Product

Liquid Vapor Feed Streams

1 596.0 -88.2 3345.9 57109.8' 62724.2²

2 596.5 -67.6 2905.8 8960.3

3 597.0 -52.5 2680.0 8520.2

4 597.5 -42.3 2439.5 8294.4

5 598.0 -32.0 8053.9 7668.0³ 2053.6⁴

'Feed to Stage 1 is 94.04 mol % vapor. ²Product removed from Stage 1, 100 mol % vapor. ³Feed to Stage 5, 98.94 mol % vapor. ⁴Product removed from Stage 5, 0 mol % vapor.

TABLE 6

SIMULATED LIQUDD VAPOR STREAM COMPOSITIONS LEAVING EACH THEORETICAL STAGE (Mole Fraction)

Nitrogen co₂ Methane Ethane Propane i-Butane n-Butane

Stage 1

Vapor 0.00231 0.000072 0.966005 0.025421 0.005277 0.000467 0.000367

Liquid 0.000359 0.000153 0.589261 0.132705 0.130329 0.033700 0.04171 1

Stage 2

Vapor 0.000640 0.000108 0.941610 0.047192 0.008898 0.000776 0.000615

Liquid 0.000085 0.000178 0.476845 0.190340 0.161 161 0.039734 0.048783

Stage 3

Vapor 0.000561 0.000115 0.921470 0.062431 0.013142 0.001 134 0.000905

Liquid 0.000069 0.000157 0.415375 0.208673 0.187549 0.044244 0.053820

Stage 4

Vapor 0.000569 0.000106 0.913713 0.064872 0.017638 0.001540 0.001229

Liquid 0.000065 0.000130 0.380377 0.191896 0.216335 0.050645 0.061013

Stage 5

Vapor 0.000583 0.000097 0.917993 0.055497 0.021253 0.002204 0.001837

Liquid 0.000060 0.000106 0.346889 0.145714 0.227598 0.062744 0.078635

TABLE 6

SIMULATED LIQUID/VAPOR STREAM COMPOSITIONS

LEAVING EACH THEORETICAL STAGE (Mole Fraction)

(CONTINUED) i-Pentane n-Pentane n-Hexane n-Heptane n-Octane Benzene n-Nonane

Stage 1

Vapor 0.000049 0.000026 0.000012 1.70E-06 1.11E-07 2.83E-06 8.51E-10

Liquid 0.015796 0.012679 0.022699 0.012625 0.003116 0.004784 0.000083

Stage 2

Vapor 0.000084 0.000046 0.000021 3.26E-06 2.23E-07 4.90E-06 1.78E-09

Liquid 0.018298 0.014662 0.026170 0.014543 0.003588 0.005516 0.000095

Stage 3

Vapor 0.000126 0.000069 0.000034 5.40E-06 3.87E-07 7.60E-06 3.21 E-09

Liquid 0.019970 0.015971 0.028414 0.015775 0.003891 0.005988 0.000103

Stage 4

Vapor 0.000171 0.000095 0.000047 7.71E-06 5.67E-07 0.000010 4.82E-09

Liquid 0.022257 0.017730 0.031314 0.017348 0.004276 0.006598 0.000114

Stage 5

Vapor 0.000273 0.000154 0.000079 0.000013 9.77E-07 0.000017 8.41E-09

Liquid 0.030295 0.024383 0.043792 0.024376 0.006019 0.009229 0.000160

Claims

THAT WHICH IS CLAIMED:

1. A process for removing and concentrating the higher molecular weight

hydrocarbon species from a methane-based gas stream comprising the steps of:

(a) condensing a minor portion ofthe methane-based gas stream thereby

producing a two-phase stream;

(b) feeding said two-phase stream into the upper section of a column;

(c) removing from the upper section of said column a heavies-depleted gas

stream;

(d) removing from the lower section of said column a heavies-rich liquid

stream;

(e) contacting via indirect heat exchange the heavies-rich liquid stream with a

methane-rich stripping gas stream thereby producing a warmed heavies-

rich stream and a cooled methane-rich stripping gas stream;

(f) feeding said cooled methane-rich stripping gas stream to the lower section

ofthe column; and

(g) contacting the two-phase stream and the cooled methane-rich stripping gas

stream in said column thereby producing the heavies-depleted gas stream

and the heavies-rich liquid stream.

2. A process according to claim 1 wherein step (a) is comprised of splitting

the methane-based gas stream into a first stream and a second stream, cooling said first stream

thereby producing a partially condensed first stream, and combining said partially condensed

first stream with the second stream thereby producing said two-phase stream.

3. A process according to claim 2 wherein the amount of liquids in said two-

phase stream is controlled by determining for the methane-based gas stream a two-phase stream

temperature corresponding to the desired liquids content at equilibrium conditions, measuring the

temperature ofthe two-phase sfream, maintaining constant the flowrate ofthe first stream and the

amount of cooling imparted to said stream, and adjusting the flowrate of said second stream

responsive to the two-phase stream temperature such that the two-phase stream temperature

approximates the calculated two-phase stream temperature.

4. A process according to claim 1 additionally comprising the step of

(h) sequentially cooling the methane-based gas stream prior to step (a) by

flowing said stream through at least one indirect heat exchange means in

contact with a first refrigerant stream thereby producing a cooled

methane-based gas stream and flowing the cooled methane-based gas

stream through at least one indirect heat exchange means in contact with a

second refrigerant stream wherein the boiling point ofthe second

refrigerant stream is less than the boiling point ofthe first refrigerating

stream thereby producing the feedstream to step (a).

5. A process according to claim 4 wherein said first refrigerant stream is

comprised in major portion of propane and said second refrigerant stream is comprised in major

portion of ethane, ethylene or a mixture thereof.

6. A process according to claim 4 further comprising:

(i) withdrawing a side stream from the methane-based gas stream at a

location downstream of one ofthe indirect heat exchange means and

employing said side stream as the methane-rich stripping gas in step (e).

7. A process according to claim 4 wherein said cooling by at least one

indirect heat exchange means in contact with a first refrigerant stream is comprised of flowing

said gas stream to be cooled through two or more indirect heat exchange means in a sequential

manner and wherein the first refrigerant to each such indirect heat exchange means has been

flashed to a progressively lower temperature and pressure in a sequentially consistent manner

and wherein said cooling by at least one indirect heat exchange means in contact with a second

refrigerant stream is comprised of flowing said gas stream to be cooled through two or more

indirect heat exchange means in a sequential manner and wherein the second refrigerant to each

indirect heat exchange means has been flashed to a progressively lower temperature and pressure

in a sequentially consistent manner

8. A process according to claim 7 wherein three indirect heat exchange

means are employed for cooling by the first refrigerant stream and two or three indirect heat

exchange means are employed for cooling by the second refrigerant stream.

9. A process according to claim 7 wherein the pressure ofthe methane-based

feed gas is 500 to 900 psia.

10. A process according to claim 7 wherein the pressure ofthe methane-based

feed gas is about 575 to about 650 psia.

11. A process according to claim 10 further comprising:

(i) withdrawing a side stream from the methane-based gas stream at a

location downstream of one ofthe indirect heat exchange means and

employing said side stream as the methane-rich stripping gas in step (e).

12. A process according to claim 1 additionally comprising:

(h) feeding the warmed heavies-rich stream of step (e) to a demethanizer

comprised of a fractionator, a reboiler and a condenser thereby producing

a heavies-rich liquid stream and a methane-rich vapor stream.

13. A process according to claim 12 wherein a major portion of the cooling

duty for the condenser is provided by the heavies-rich liquid stream produced by step (d) or step

(e).

14. A process according to claim 12 wherein a major portion ofthe cooling

duty for the condenser is provided by flowing through an indirect heat exchange means in

contact with the heavies-rich liquid stream of step (d) and the resulting treated heavies-rich liquid

stream becomes the heavies-bearing feedstream to step(e).

15. A process according to claim 13 wherein the cooling duty is provided by

splitting the overhead vapor stream into a first vapor stream and a second vapor stream, cooling

and partially condensing said first stream via indirect heat exchange with the heavies rich liquid

stream of step (d ) thereby producing a cooled, partially condensed first stream, combining said

first stream and said second stream, feeding said combined stream to a gas-liquid separator from

which is produced the reflux stream to the fractionating column and the methane-rich vapor

stream.

16. A process according to claim 15 wherein the flowrate ofthe reflux stream

is controlled by calculating for the overhead vapor stream a two-phase stream temperature

corresponding to the desired liquids content at equilibrium conditions, measuring the temperature

ofthe two-phase stream, maintaining constant the flowrate ofthe first stream and the amount of

cooling imparted to said stream, and adjusting the flowrate of said second stream responsive to the two-phase stream temperature such that the calculated two-phase stream temperature is

approached .

17. A process according to claim 13 additionally comprising between steps

(d) and (e) the additional step of :

(i) flashing the heavies-rich liquid stream to a lower pressure thereby further

decreasing the temperature of said stream.

18. A process according to claim 17 additionally comprising the step of

(j) condensing the heavies depleted gas stream thereby producing a liquefied

natural gas stream.

19. A process according to claim 18 wherein said condensing is comprised of

flowing the heavies depleted gas stream through an indirect heat exchange means cooled by said

second refrigerant stream.

20. A process according to claim 19 wherein the pressure ofthe methane-

based gas stream is 500 to 900 psia.

21. A process according to claim 20 additionally comprising the steps of

(k) flashing in one or more steps the liquefied product of step (j) to

approximately atmospheric pressure thereby producing an LNG product

stream and one or more methane vapor streams;

(1) compressing a majority ofthe vapor streams of step (k) to a pressure of

500 to 900 psia,

(m) cooling said compressed vapor stream of step (1); and (n) combining the resulting cooled stream with the methane-based gas stream

fed to step (a) or the resulting product from one ofthe indirect heat

exchange means of step (h).

22. A process according to claim 21 wherein the methane-rich vapor stream of

step (h) is combined with one ofthe vapor streams of step (k) prior to step (1).

23. A process according to claim 21 wherein the pressure ofthe

methane-based feed gas and the gas stream from step (1) is about 575 to about 650 psia.

24. A process according to claim 1 wherein the column provides two to

fifteen theoretical stages of gas-liquid contacting.

25. A process according to claim 1 wherein the column provides three to ten

theoretical stages of gas-liquid contacting.

26. A process according to claim 23 wherein the column provides two to

fifteen theoretical stages of gas-liquid contacting.

27. A process according to claim 23 wherein the column provides three to ten

theoretical stages of gas-liquid contacting.

28. A process for removing benzene and other aromatics from a

methane-based gas sfream comprising the steps of:

(a) condensing a minor portion of the methane-based gas stream thereby

producing a two-phase stream;

(b) feeding said two-phase stream into the upper section of a column;

(c) removing from the upper section of said column a

benzene/aromatic-depleted gas stream; (d) removing from the lower section of said column a benzene/aromatic-rich

liquid stream;

(e) contacting via indirect heat exchange the benzene/aromatic-rich liquid

stream with a methane-rich stripping gas stream thereby producing a

warmed benzene/aromatic-rich stream and a cooled methane-rich stripping

gas stream;

(f) feeding said cooled methane-rich stripping gas sfream to the lower section

ofthe column; and

(g) contacting the two-phase stream and the cooled methane-rich stripping gas

sfream in said column thereby producing the benzene/aromatic-depleted

gas stream and the benzene/aromatic-rich liquid stream.

29. A process according to claim 28 wherein step (a) is comprised of splitting

the methane-based gas stream into a first stream and a second stream, cooling said first stream

thereby producing a partially condensed first stream, and combining said partially condensed

first sfream with the second stream thereby producing said two-phase stream.

30. A process according to claim 29 wherein the amount of liquids in said

two-phase stream is controlled by determining for the methane-based gas stream a two-phase

stream temperature corresponding to the desired liquids content at equilibrium conditions,

measuring the temperature ofthe two-phase stream, maintaining constant the flowrate ofthe first

stream and the amount of cooling imparted to said stream, and adjusting the flowrate of said

second stream responsive to the two-phase stream temperature such that the two-phase stream

temperature approximates the calculated two-phase stream temperature.

31. A process according to claim 28 additionally comprising the step of

(h) sequentially cooling the methane-based gas stream prior to step (a) by

flowing said stream through at least one indirect heat exchange means in

contact with a first refrigerant stream thereby producing a cooled

methane-based gas stream and flowing the cooled methane-based gas

stream through at least one indirect heat exchange means in contact with a

second refrigerant stream where the boiling point ofthe second refrigerant

stream is less than the boiling point ofthe first refrigerating stream

thereby producing the feedstream to step (a).

32. A process according to claim 31 wherein said first refrigerant stream is

comprised in major portion of propane and said second refrigerant stream is comprised in major

portion of ethane, ethylene or a mixture thereof.

33. A process according to claim 31 further comprising:

(i) withdrawing a side stream from the methane-based gas stream at a

location downstream of one of the indirect heat exchange means and

employing said side stream as the methane-rich stripping gas in step (e).

34. A process according to claim 31 wherein said cooling by at least one

indirect heat exchange means in contact with a first refrigerant stream is comprised of flowing

said gas stream to be cooled through two or more indirect heat exchange means in a sequential

manner and wherein the first refrigerant to each such indirect heat exchange means has been

flashed to a progressively lower temperature and pressure in a sequentially consistent manner

and wherein said cooling by at least one indirect heat exchange means in contact with a second

refrigerant stream is comprised of flowing said gas stream to be cooled through two or more indirect heat exchange means in a sequential manner and wherein the second refrigerant to each

indirect heat exchange means has been flashed to a progressively lower temperature and pressure

in a sequentially consistent manner

35. A process according to claim 34 wherein three indirect heat exchange

means are employed for cooling by the first refrigerant stream and two or three indirect heat

exchange means are employed for cooling by the second refrigerant stream.

36. A process according to claim 34 wherein the pressure ofthe

methane-based feed gas is 500 to 900 psia.

37. A process according to claim 34 wherein the pressure ofthe methane-

based feed gas is about 575 to about 650 psia.

38. A process according to claim 37 further comprising:

(i) withdrawing a side stream from the methane-based gas stream at a

location downstream of one ofthe indirect heat exchange means and

employing said side stream as the methane-rich stripping gas in step (e).

39. A process according to claim 28 additionally comprising:

(h) feeding the warmed benzene/aromatic-rich stream of step (e) to a

demethanizer comprised ofa fractionator column, a reboiler and a

condenser thereby producing a benzene/aromatic-rich liquid stream and a

methane-rich vapor sfream.

40. A process according to claim 39 wherein a major portion of the cooling

duty for the condenser is provided by the benzene/aromatic-rich liquid stream produced by step

(d) or step (e).

41. A process according to claim 39 wherein a major portion ofthe cooling

duty for the condenser is provided by flowing through an indirect heat exchange means in

contact with the benzene/aromatic-rich liquid sfream of step (d) and the resulting treated

benzene/aromatic-rich liquid stream becomes the benzene/aromatic-bearing feedstream to

step(e).

42. A process according to claim 40 wherein the cooling duty is provided by

splitting the overhead vapor stream into a first vapor stream and a second vapor stream, cooling

and partially condensing said first stream via indirect heat exchange with the benzene/aromatic

rich liquid stream of step (d) thereby producing a cooled, partially condensed first stream,

combining said first stream and said second stream, feeding said combined stream to a gas-liquid

separator from which is produced the reflux stream to the fractionating column and the methane-

rich vapor stream.

43. A process according to claim 42 wherein the flowrate ofthe reflux stream

is controlled by calculating for the overhead vapor stream a two-phase stream temperature

corresponding to the desired liquids content at equilibrium conditions, measuring the temperature

ofthe two-phase stream, maintaining constant the flowrate ofthe first stream and the amount of

cooling imparted to said stream, and adjusting the flowrate of said second sfream responsive to

the two-phase stream temperature such that the calculated two-phase stream temperature is

approached.

44. A process according to claim 40 additionally comprising between steps

(d) and (e) the additional step of :

(i ) flashing the benzene/aromatic-rich liquid stream to a lower pressure

thereby further decreasing the temperature of said stream.

45. A process according to claim 44 additionally comprising the step of

(j) condensing the benzene/aromatic depleted gas stream thereby producing a

liquefied natural gas stream.

46. A process according to claim 45 wherein said condensing is comprised of

flowing the benzene/aromatic depleted gas stream through an indirect heat exchange means

cooled by said second refrigerant stream.

47. A process according to claim 46 wherein the pressure ofthe methane-

based gas stream is 500 to 900 psia.

48. A process according to claim 47 additionally comprising the steps of

(k) flashing in one or more steps the liquefied product of step (j) to

approximately atmospheric pressure thereby producing an LNG product

stream and one or more methane vapor streams;

(1) compressing a majority ofthe vapor streams of step (k) to a pressure of

500 to 900 psia,

(m) cooling said compressed vapor sfream of step (1); and

(n) combining the resulting cooled stream with d e methane-based gas stream

fed to step (a) or the resulting product from one ofthe indirect heat

exchange means of step (h).

49. A process according to claim 48 wherein the methane-rich vapor stream of

step (h) is combined with one ofthe vapor streams of step (k) prior to step (o).

50. A process according to claim 48 wherein the pressure of the

methane-based feed gas and the gas stream from step (1) is about 575 to about 650 psia.

51. A process according to claim 28 wherein the column provides two to

fifteen theoretical stages of gas-liquid contacting.

52. A process according to claim 28 wherein the column provides three to ten

theoretical stages of gas-liquid contacting.

53. A process according to claim 50 wherein the column provides two to

fifteen theoretical stages of gas-liquid contacting.

54. A process according to claim 50 wherein the column provides three to ten

theoretical stages of gas-liquid contacting.

55. A apparatus comprising:

(a) a condenser;

(b) a column;

(c) a heat exchanger providing for indirect heat exchange between two fluids;

(d) a conduit between said condenser and the upper section ofthe column for

flow of a two-phase stream to the column ;

(e) a second conduit connected to the upper section of the column for the

removal of a vapor stream from the column;

(f) a conduit between said column and heat exchanger for flow of a cooled

gas stream from the heat exchanger;

(g) a conduit between said column and said heat exchanger for flow of a

liquid stream from the column;

(h) a conduit connected to the heat exchanger for flow of a warmed liquid

stream from the heat exchanger; and (i) a conduit connect to the heat exchanger for flow ofa gas stream to the

heat exchange.

56. A apparatus according to claim 55 additionally comprised of a

(j) a first conduit;

(k) a splitting means connected to the first conduit;

(1) a second conduit and a third conduit connected to said splitting means

where said second conduit is connected to the condenser;

(m) a control valve connected at the inlet side to the second conduit,

(n) a conduit connected to the outlet side of said control valve;

(o) a junction or combining means connected to said conduit of element (n)

and the conduit of element (d) prior to connection with the column;

(p) a temperature sensing means with sensing element situated in conduit of

element (d) between said junction means and connection with the column;

and

(q) a control means operably attached to control valve of element (m) and

operably responsive to input received from the temperature sensing device

of element (p) and a temperature setpoint.

57. An apparatus according to claim 55 additionally comprising of

(j) a pressure reduction means situated in conduit (g).

58. An apparatus according to claim 55 wherein said column contains 2 to 12

theoretical stages.

59. An apparatus according to claim 55 additional comprising one or more

indirect heat exchange means situated in a sequential manner, conduits between each heat exchange means for the sequential flow of a common fluid through the heat exchangers

whereupon the last conduit is connected to the condenser of element (a), conduits to and from

each heat exchanger providing for the flow of a refrigerating agent to each heat exchanger and

wherein the conduit of element (i) is in flow communication with one ofthe above conduits for

flow of a common fluid between heat exchangers.

60. An apparatus according to claim 59 wherein propane is employed as the

refrigerating agent in at least two ofthe heat exchange means; and ethane, ethylene or a mixture thereof is employed as the refrigerating agent in at least two heat exchange means.

61. A apparatus according to claim 55 additionally comprising:

(j) a fractionation column;

(k) a reboiler;

(1) a condenser;

(m) an overhead conduit connecting the upper section ofthe column to the

condenser for removal ofthe overhead vapor, a reflux conduit connected

the condenser to the column for the return of the reflux fluid, a vapor

product conduit connected to the condenser for removal of uncondensed

vapors;

(n) a bottoms conduit connecting the lower section of the column to the

reboiler, a vapor conduit for returning stripping vapor to the column, and a

bottoms product line connected to the reboiler for removal of unvaporized

product from the reboiler; and

wherein the conduit of element (h) is connected to the fractionation column at a point between

the top and the bottom theoretical stages.

62. A apparatus according to claim 61 wherein the condenser of element (1) is

comprised of an indirect heat exchange means and coolant to such means is provided by a

junction connecting the cooling side ofthe indirect heat exchange means to the conduit of

element (g).

63. An apparatus according to claim 61 additionally comprising

(o) a pressure reduction means situated in conduit (g) and

wherein the condenser of element (k) is comprised of an indirect heat

exchange means and said coolant to such means is provided by a junction

connecting the cooling side ofthe indirect heat exchange means to the

conduit of element (g) downstream of pressure reduction means (o).

64. An apparatus according to claim 61 additionally comprising a

(o) a conduit connected to condenser of element (a),

(p) a compressor connected at the inlet port to the vapor conduit line of

element (m); and

(q) a conduit connecting the outlet port of said compressor element (p) to the

conduit of element (o).

65. A apparatus according to claim 59 additionally comprising:

(j) a fractionation column;

(k) a reboiler;

(1) a condenser;

(m) an overhead conduit connecting the upper section ofthe column to the

condenser for removal ofthe overhead vapor, a reflux conduit connected

the condenser to the column for the return ofthe reflux fluid, a vapor product conduit connected to the condenser for removal of uncondensed

vapors;

(n) a bottoms conduit connecting the lower section ofthe column to the

reboiler, a vapor conduit for returning stripping vapor to the column, and a

bottoms product line connected to the reboiler for removal of unvaporized

product from the reboiler; and

wherein the conduit of element (h) is connected to the fractionation

column at a midpoint location.

66. An apparatus according to claim 65 additionally comprising a

(o) a compressor connected at the inlet port to the vapor conduit line of

element (m) and

(p) conduit connecting the outlet port of said compressor to one of the

common flow conduits of claim 59.

67. Apparatus comprising:

(a) a cryogenic separation column for partially condensing a feed gas stream in

an LNG recovery process;

(b) means for withdrawing a liquid condensate stream from said cryogenic

separation column.

(c) a heat exchanger associated with said cryogenic separation column;

(d) means for passing said liquid condensate stream through said heat exchanger;

(e) means for passing a warm dry gas stream through said heat exchanger and

thereafter to said cryogenic separation column, wherein said warm dry gas stream is cooled by

indirect heat exchange with said liquid condensate stream in said heat exchanger; (f) a bypass conduit having a first control valve operably located therein for

bypassing said warm dry gas stream around said heat exchanger;

(g) means for establishing a first signal representative ofthe actual temperature of

said warm dry gas stream exiting said heat exchanger;

(h) means for establishing a second signal representative ofthe actual temperature

of said liquid condensate stream entering said heat exchanger;

(i) means for dividing said first signal by said second signal to establish a third

signal representative ofthe ratio of said first signal and said second signal;

(j) means for establishing a fourth signal representative ofa desired value for the

ratio represented by said third signal;

(k) means for comparing said third signal and said fourth signal and establishing a

fifth signal which is responsive to the difference of said third signal and said fourth signal,

wherein said fifth signal is scaled to be representative of the position of said first control valve

required to maintain the actual ratio represented by said third signal substantially equal to the

desired ratio represented by said fourth signal; and

(m) means for manipulating said first control valve in said bypass conduit in

response to said fifth signal.

68. Apparatus in accordance with claim 67, additionally comprising:

means for establishing a sixth signal scaled to be representative ofthe flow rate of

said liquid condensate stream required to maintain a desired liquid level in said cryogenic

separation column; and

means for controlling the flow rate of said liquid condensate stream responsive to

said sixth signal.

69. Apparatus in accordance with claim 68, additionally comprising:

a second control valve operably located so as to control flow of said warm dry gas

stream; and

means for manipulating said second control valve responsive to a temperature selected from the pair of temperatures consisting of:

i. the actual temperature of said warm dry gas stream exiting said heat

exchanger; and

ii. the actual temperature of said liquid condensate stream exiting said heat exchanger.

70. Apparatus in accordance with claim 69, wherein said means for

manipulating said second control valve comprises:

means for establishing a seventh signal representative ofthe actual temperature of

said liquid condensate stream exiting said heat exchanger;

means for establishing an eighth signal representative ofthe desired temperature

of said liquid condensate stream exiting said heat exchanger;

means for comparing said seventh signal and said eighth signal to establish a

ninth signal responsive to the difference of said seventh signal and said eighth signal, wherein

said ninth signal is scaled to be representative ofthe position of said second control valve

required to maintain the actual temperature of said liquid condensate stream exiting said heat

exchanger represented by said seventh signal substantially equal to the desired temperature

represented by said eighth signal; means for establishing a tenth signal representative of the desired temperature of

said warm dry gas stream exiting said heat exchanger represented by said second signal; means for comparing said second signal and said tenth signal to establish an

eleventh signal responsive to the difference between said second signal and said tenth signal,

wherein said eleventh signal is scaled to be representative ofthe position of said second control

valve required to maintain the actual temperature of said warm dry gas stream exiting said heat

exchanger substantially equal to the desired value represented by said tenth signal;

means for establishing a twelfth signal selected as the one of said ninth signal and

said eleventh signal having the higher value; and

means for manipulating said second control valve responsive to said twelfth

signal.

71. A method for controlling temperature in a heat exchanger equipped with a

bypass conduit having a first control valve operatively connected therein, said heat exchanger

being associated with a cryogenic separation column that removes a benzene contaminant from a feed stream in and LNG recovery process, said method comprising:

withdrawing a liquid condensate stream at a cryogenic temperature from said

cryogenic separation column;

passing said liquid condensate stream through said heat exchanger;

passing a warm dry gas stream through said heat exchanger and thereafter

infroducing said warm dry gas stream into said cryogenic separation column, wherein said warm

dry gas stream is cooled by indirect heat exchange with said liquid condensate stream in said heat exchanger;

establishing a first signal representative ofthe actual temperature of said warm

dry gas stream exiting said heat exchanger; establishing a second signal representative ofthe actual temperature of said liquid

condensate stream entering said heat exchanger;

dividing said first signal by said second signal to establish a third signal

representative ofthe ratio of said first signal and said second signal;

establishing a fourth signal representative of a desired value for said third signal;

comparing said third signal and said fourth signal and establishing a fifth signal which is responsive to the difference between said third signal and said fourth signal, wherein

said fifth signal is scaled to be representative ofthe position of said first control valve required to

maintain the actual ratio represented by said third signal substantially equal to the desired ratio

represented by said fourth signal; and

manipulating said first confrol valve in said bypass conduit in response to said

fifth signal.

72. A method in accordance with claim 71 additionally comprising the

following steps:

establishing a sixth signal scaled to be representative of the flow rate of said

liquid condensate steam required to maintain a desired liquid level in said cryogenic separation

column; and

controlling the flow rate of said liquid condensate stream responsive to said sixth

signal.

73. A method in accordance with claim 71 , wherein a second control valve is

operably located so as to control flow rate of said warm dry gas stream, said method additionally

comprising the following steps: manipulating said second control valve responsive to a temperature selected from

the pair of temperatures consisting of:

i) the actual temperature of said warm dry gas stream exiting said heat

exchanger; and

ii) the actual temperature of said liquid condensate stream exiting said

heat exchanger.

74. A method in accordance with claim 73, wherein said step of manipulating

said second control valve comprises:

establishing a seventh signal representative ofthe actual temperature of said

liquid condensate stream exiting said heat exchanger;

establishing an eighth signal representative ofthe desired temperature of said

liquid condensate stream exiting said heat exchanger;

comparing said seventh signal and said eighth signal to establish a ninth signal

responsive to the difference between said seventh signal and said eighth signal, wherein said

ninth signal is scaled to be representative of the position of said second control valve required to

maintain the actual temperature of said liquid condensate stream exiting said heat exchanger

represented by said seventh signal substantially equal to the desired temperature represented by

said eighth signal;

establishing a tenth signal representative ofthe desired temperature of said warm dry gas stream exiting said heat exchanger represented by said second signal;

comparing said second signal and said tenth signal to establish an eleventh signal

responsive to the difference between said second signal and said tenth signal, wherein said

eleventh signal is scaled to be representative ofthe position of said second control valve required to maintain the actual temperature of said warm dry gas stream exiting said heat exchanger

substantially equal to the desired value represented by said tenth signal;

establishing a twelfth signal selected as the one of said ninth signal and said

eleventh signal having the higher value; and

manipulating said second control valve responsive to said twelfth signal.

75. A method in accordance with claim 67, wherein said LNG recovery

process is a cascade refrigeration process employing three different refrigerants.