Patent Description:
The written disclosure herein describes illustrative embodiments that are non-limiting and non-exhaustive. Reference is made to certain of such illustrative embodiments that are depicted in the figures, in which:.

Methods of cryogenically purifying process streams and systems related thereto are disclosed herein. It will be readily understood that the embodiments as generally described below and as illustrated in the examples and Figures could be modified in a wide variety of ways. Thus, the following more detailed description of various embodiments, as described below and represented in the examples and Figures, is not intended to limit the scope of the invention, but is merely representative of various embodiments. <CIT> discloses a method of separating carbon dioxide from a process stream according to the preamble of independent claim <NUM>.

The phrase "in communication with" and the term "connecting" refer to any form of interaction between two or more entities, including mechanical, electrical, magnetic, electromagnetic, fluid, and thermal interaction. Two entities may interact with each other even though they are not in direct contact with each other. For example, two entities may interact with each other through an intermediate entity.

Methods of cryogenically purifying process streams may comprise methods of separating carbon dioxide from a process stream and condensing a primary component of the process stream. Alternatively, methods of cryogenically purifying process streams may comprise methods of separating carbon dioxide from a process stream without condensing a primary component of the process stream. As defined in claim <NUM>, the present invention as claimed is directed towards a method of separating carbon dioxide from a process stream comprising (among other steps recited in claim <NUM>) condensing the primary component and solidifying the carbon dioxide.

In some examples of methods of separating carbon dioxide from a process stream and condensing a primary component of the process stream that do not fall within the scope of the appended claims, the methods comprise providing a process stream comprising gaseous carbon dioxide and further comprising a primary component, wherein the process stream is at a first temperature and a first pressure. The methods further comprise cooling the process stream to at or below the condensation temperature of the primary component in the process stream. The methods may further comprise separating any gases from the process stream that did not condense during cooling the process stream to at or below the condensation temperature of the primary component to form a first separated gaseous stream. The methods further comprise cooling the process stream further to at or below the frost point of carbon dioxide in the process stream, thereby forming solid carbon dioxide, and separating physically solid carbon dioxide from the process stream to form a separated solid carbon dioxide slurry stream and a liquid primary component stream.

In some examples of separating carbon dioxide from a process stream without condensing a primary component of the process stream that do not fall within the scope of the appended claims, the methods comprise providing a process stream comprising gaseous carbon dioxide and further comprising a primary component, wherein the process stream is at a first temperature and a first pressure. The methods further comprise reducing the temperature of the process stream to a temperature at or below the frost point of carbon dioxide in the process stream, by directly contacting the process stream with a colder contact liquid, thereby forming solid carbon dioxide in the contact liquid and forming a gaseous purified primary component stream. The methods may further comprise separating physically solid carbon dioxide from the contact liquid to form a solid carbon dioxide slurry stream from a purified contact liquid stream.

Systems (not falling within the scope of the appended claims) for cryogenically purifying process streams may comprise systems for separating carbon dioxide from a process stream and condensing a primary component of the process stream. Systems (not falling within the scope of the appended claims) for cryogenically purifying process streams may comprise systems for separating carbon dioxide from a process stream without condensing a primary component of the process stream.

In some examples of systems (not falling within the scope of the appended claims) for separating carbon dioxide from a process stream and condensing a primary component of the process stream, the systems comprise a desublimation heat exchanger configured to receive a process stream comprising gaseous carbon dioxide and a primary component with a condensation temperature above the frost point of carbon dioxide at the pressure of the process stream, the desublimation heat exchanger further configured to cool the process stream to at or below the frost point of carbon dioxide in the process stream. The systems may further comprise a solid-liquid separator configured to physically separate a solid carbon dioxide slurry stream from a liquid purified primary component stream.

In some examples of systems (not falling within the scope of the appended claims) for separating carbon dioxide from a process stream without condensing a primary component of the process stream, the systems comprise a desublimation heat exchanger configured to receive a process stream comprising gaseous carbon dioxide and a primary component, configured to receive a colder contact liquid stream, configured to directly contact the process stream with the colder contact liquid stream and cool the process stream to at or below the frost point of carbon dioxide in the process stream, configured to produce a gaseous purified primary component stream, and configured to produce a solids-containing contact liquid stream. The systems may further comprise a solid-liquid separator configured to receive the solids-containing contact liquid stream and physically separate a solid carbon dioxide slurry stream from a purified contact liquid stream.

The methods and systems disclosed herein (not all of which fall within the scope of the appended claims but are provided for exemplary purposes) may be used to remove carbon dioxide ("CO<NUM>") and condensable or absorbable liquids from a process stream. For example, the methods and systems may be used for removing CO<NUM>, natural gas liquids, and other components that condense and/or absorb under the specified operating conditions from a raw natural gas stream. In another example, the methods and systems may be used for treating syngas/producer gas stream for CO<NUM> and other condensable/absorbable components. In both cases, the process produces a treated stream with less CO<NUM> and condensable liquids. Additionally, a liquid CO<NUM> stream and a separated liquids stream may be produced as well.

The methods and systems (not all of which fall within the scope of the appended claims but are provided for exemplary purposes) may be used to treat both high-pressure streams, meaning streams at pressures and temperatures under which the bulk of the stream condenses to form a liquid in some portion of the process, and low-pressure streams, meaning streams at pressures and temperatures under which the bulk of the treated stream remains a gas but that may nevertheless be at greater than ambient pressure. Process descriptions for these two exemplesappear in separate descriptions below.

LNG quality requirements are <NUM>% and <NUM>% CO<NUM> in Canada and the European Union, respectively (Grunvald, A. Izotov and V. Nemov (<NUM>). Gas Quality Requirements as a Factor of Successful LNG Projects Implementation. International Gas Union Research Conference, Paris. The United States does not currently have LNG quality requirements, but re-vaporized LNG must meet natural gas pipeline standards of <NUM>-<NUM>% CO<NUM> (Gas Processors Suppliers Association (<NUM>). GPSA Engineering Data Book. Even with these standards, in practice natural gas is pretreated for CO<NUM> removal down to <NUM> ppm before liquefaction. The reason for CO<NUM> removal down to <NUM> ppm stems from concerns with degradation of operability due to potential CO<NUM> freezing during natural gas liquefaction. The solubility of CO<NUM> in the final LNG product at normal conditions is higher than <NUM> ppm (<NPL>.

An exemplary process configuration for the condensing or high-pressure process appears in <FIG> with major equipment described in Table <NUM>. The discussion focuses on natural gas processing; however, the technology applies to any process in which a major portion of the stream to be treated condenses as a liquid (at the pressures of the stream and at a temperature higher than the frost point of CO<NUM> in the stream). The dashed, solid, and dot-dashed lines in the Figure represent streams that are primarily gaseous, liquid, and solid, respectively. Multiple phases may exist in many of the streams, with the type of line representing the dominant phase.

<FIG> includes three subprocesses, namely a natural gas treating loop, a melting loop, and a low-temperature cooling loop. Only the first of these is required to implement the reducing gas cryogenic carbon capture process, with the possible requirement of supplemental refrigeration of traditional types. Descriptions of all three exemplary subprocesses appear separately below. The main processes are removal of natural gas liquids and CO<NUM> from the raw natural gas stream. These processes occur in the natural gas treating loop and are described first.

The natural gas treating loop involves the major pieces of equipment identified in Table <NUM> and the associated connecting streams, minor equipment, and controls, listed in the order that the incoming raw natural gas stream encounters them.

The exemplary process begins with raw natural gas (the "process stream" at the "first pressure" of this example). This gas cools to ambient or slightly below ambient temperature using cooling water or other resources (the "first temperature" of this example) and is then dried using conventional techniques that are not illustrated. The raw, dry natural gas feed stream first cools in heat exchanger E-<NUM> (the "first heat exchanger" of this example) to temperatures above and near the melting point of pure CO<NUM> at the pressure of the system (the "second temperature" of this example), which in most cases will be near -<NUM>. "Above and near" may comprise less than about <NUM>° C above, less than about <NUM>° C above, or less than about <NUM>° C above. The first heat exchanger condenses some of the natural gas liquids present in the process stream as the stream cools. These natural gas liquids separate (the "first separated liquids stream" of this example) from the gaseous bulk of the flow in E-<NUM> and combine with other natural gas liquid streams originating from other portions of the process to form stream P-<NUM>. Stream P-<NUM> returns counter-currently through the heat exchanger to return to near the first temperature as it helps cool the process stream. A liquid pump, E-<NUM>, included as part of E-<NUM>, raises the pressure of the stream P-<NUM> such that it can mix with stream P-<NUM> to form stream P-<NUM>. Stream P-<NUM> may be at sufficient pressure that none of it, or only the lightest components of it, will vaporize as it warms back to the first temperature in E-<NUM>. In other embodiments, some or all of the stream will vaporize.

In an alternative embodiment, the liquids collected in E-<NUM> flow in a separate stream through E-<NUM> and rejoin the vapors in stream P-<NUM>, with or without any liquids collected in E-<NUM>. In this alternative embodiment, any liquids condensed may be combined with the condensed primary component formed in E-<NUM>.

The bulk of the natural gas stream (the "process stream") flows as a gas from E-<NUM> into a second heat exchanger, E-<NUM>, that drops the temperature to above and near the condensation point of the natural gas at the pressure of the system (the "third temperature" of this example), which is typically about -<NUM>. Some trace impurities, such as Hg, will largely condense with the liquids in E-<NUM> while others, such as H<NUM>S, may not condense until later in the process, depending on their concentrations in the raw gas. Other trace components may never condense, such as noble gases like helium and argon. Additional natural gas liquids condense as the stream cools, and E-<NUM> separates these natural gas liquids (the "second separated liquids stream" of this example) from the gaseous stream and introduces them to the natural gas liquids stream P-<NUM> in a manner similar to E-<NUM>, namely, by pumping them to the line pressure. They return through E-<NUM> and help cool the incoming natural gas stream as they warm.

The bulk of the natural gas stream (the "process stream") leaves E-<NUM> as a vapor and enters a third heat exchanger E-<NUM>, where the methane condenses to form some or mostly liquid as stream P-<NUM>. The now-liquid natural gas stream exits the heat exchanger E-<NUM>. Depending on the initial gas composition, it is possible that stream P-<NUM> will contain CO<NUM> or other gas-phase species. These are separated in stream P-<NUM> (the "third separator" of this example) to join the solid CO<NUM> stream at this point. Non-condensed trace components, such as noble gases, could also be separated at this point from stream P-<NUM>. Alternatively, non-condensed gases may be separated slightly later in the process in versions that do not have a melting loop, such as discussed below. The liquid natural gas (the "process stream") passes through a pressure regulating device shown as valve V-<NUM> in the diagram but which could also be a cryogenic turbine.

The natural gas stream which is now line P-<NUM> passes into a solids-forming heat exchanger (the "desublimation heat exchanger" of this example) E-<NUM>, similar to those disclosed in <CIT> or <CIT>, or of other suitable type. The desublimation heat exchanger may be a direct-contact heat exchanger and configured to form solids in the heat exchanger without fouling or plugging and may integrate heat transfer to minimize both pressure drop and energy consumption. The heat exchanger forms solid CO<NUM> in the stream as the stream cools to its lowest temperature, which is dictated by the amount of CO<NUM> and natural gas liquids removal required by the process. Depending on the operating pressure and composition, some vapor or gases may be present in E-<NUM>, P-<NUM>, and possibly P-<NUM>. The solids-forming heat exchanger may be staged (e.g., one or more additional desublimation heat exchangers) to maximize the process efficiency.

The solid-liquid slurry, P-<NUM>, that forms in the solids-forming heat exchanger, E-<NUM>, flows into a solids-liquid separation device, E-<NUM>. The solids leave the separator, E-<NUM>, as a slurry in P-<NUM>, and combines with the residual light gases in stream P-<NUM>, if there are any. The slurry in P-<NUM> may have as little residual liquid as possible - as much as is needed for transport via suitable auger, pump, or other conveyance, E-<NUM>, back through heat exchanger E-<NUM>. The slurry in P-<NUM> helps cool the process stream as it warms. The liquids phase (the "liquid purified primary component stream" of this example), P-<NUM>, which contains most of the now-refined natural gas in a liquid phase, exits the solids-liquid separator, E-<NUM>, and flows in line P-<NUM> through a pressure control valve, V-<NUM>, and then as line P-<NUM> back into heat exchanger, E-<NUM>, where it vaporizes and warms as it helps to cool the incoming stream. The pressure of the returning natural gas liquid stream, P-<NUM>, is regulated by valve V-<NUM> or conceivably a turbine to form stream P-<NUM> to optimize the performance of the system, specifically to provide optimal temperature profile matching in E-<NUM>.

The contacting liquid in the solids-forming heat exchanger forms from the condensates of the natural gas stream, which includes the methane in this example. It will typically be saturated with CO<NUM> by the time it exits the heat exchanger at its low-temperature limit. In the case that there is insufficient condensed methane to satisfy the demand for a contacting liquid, a portion of P-<NUM> could be cooled and then recirculated to E-<NUM> to maintain adequate liquid to operate the desublimation heat exchanger.

Returning to <FIG>, the natural gas stream, P-<NUM>, exits E-<NUM> as a gas with reduced natural gas liquids and CO<NUM> contents, as well as reduced contents of other condensable gases and trace contaminants (H<NUM>S, Hg), and flows through heat exchangers E-<NUM> and E-<NUM>, cooling the incoming streams as it warms. The natural gas eventually returns to near the first temperature as a refined product. The solid CO<NUM> stream, P-<NUM>, exits E-<NUM> after helping to cool the incoming process stream, P-<NUM>, and continues to warm as it cools the incoming gas process stream, P-<NUM>, in heat exchanger E-<NUM>.

Alternatively, the liquefied natural gas, such as in P-<NUM>, may not be warmed and vaporized as illustrated, but instead may be transported to an LNG distribution network, with or without additional processing.

Returning to <FIG>, the solid CO<NUM> stream exits heat exchanger E-<NUM> near its melting point and flows to a solids-liquids separation device, E-<NUM>. Routing the solids streams P-<NUM> and P-<NUM> through the heat exchangers E-<NUM> and E-<NUM> may not be practical, and the solids may go directly from E-<NUM> to E-<NUM>. The solids-liquids separator, E-<NUM>, removes residual liquids from the stream and returns these to the natural gas liquids stream as stream P-<NUM>. The remaining solids flow as stream P-<NUM> to the melting heat exchanger, E-<NUM>, where they melt as they condense stream P-<NUM>. Stream P-<NUM> is described as part of the melting loop below.

The liquid CO<NUM> stream, P-<NUM>, exiting the melting heat exchanger, E-<NUM>, is further pressurized in pump E-<NUM> and then flows through heat exchanger E-<NUM> to help cool the incoming raw natural gas feed (the "process stream") as it returns to near the first temperature as a nearly pure, liquid CO<NUM> stream. The final state of the CO<NUM> depends on the incoming stream temperature and overall pressure. The state can be liquid, supercritical, or gas.

The process pressure after the expansion valve/turbine, V-<NUM>, determines several operating conditions of the process. The example illustrated in <FIG> assumes little to no pressure drop in V-<NUM>. If the pressure drops sufficiently that methane forms some vapor, the methane in the system becomes an auto-refrigerant and can perform some or all of the cooling.

If, in addition, the amount of CO<NUM> is low or the heat recovery from the CO<NUM> is unimportant, a simple version of the process (not falling within the scope of the appended claims but provided for exemplary purposes) as illustrated in <FIG> may be used. This version of the process would usually be able to reduce CO<NUM> content to pipeline specifications (< <NUM>%) and/or to LNG specifications (< <NUM>%). In this case, the solids separator, E-<NUM>, produces a fluid stream that contains both gas and liquid, P-<NUM>, but the remainder of the process is similar to that of <FIG>. The process may need supplemental cooling at certain locations, as would be provided by refrigeration loops commonly known to process engineers.

If maintaining pressure is important or the pressure drop in V-<NUM> is insufficient for a portion of the stream to vaporize, solid CO<NUM> formation will require additional cooling. Some non-limiting examples of possible cooling mechanisms are discussed in the next section.

Depending on the composition of the raw natural gas (the "process stream") and the relative importance of operating costs, energy efficiency, and capital costs, the steps described in both of the embodiments above may suffice to perform the separations, possibly augmented by refrigerant loops to supplement cooling where needed. The process diagram in <FIG> illustrates additional innovations that increase energy efficiency and decrease overall complexity and cost, as discussed next.

Some optional cooling support for the melting heat exchanger, E-<NUM>, and the solids-forming heat exchanger, E-<NUM>, are discussed below.

The melting loop uses a nearly closed-loop refrigeration system to melt the solid CO<NUM> and transfer the heat to the coldest point in the overall system. It involves the major pieces of equipment identified in Table <NUM> and the associated connecting streams and controls.

The cooling loop compresses the refrigerant in compressor E-<NUM> and cools it in a heat exchanger E-<NUM> to or near the available heat rejection temperature of the process, which will usually be near ambient temperature. Many embodiments may use staged compressors and heat exchangers to maximize the efficiency of this process, effectively combining E-<NUM> and E-<NUM> in several stages in a single unit. The pressurized refrigerant in P-<NUM> cools to near the melting point of CO<NUM> in heat exchanger E-<NUM> and then flows through P-<NUM> and condenses in the melting heat exchanger, E-<NUM>, as it melts the solid CO<NUM>. The resulting liquid refrigerant stream, P-<NUM>, flows through a second recuperating heat exchanger, E-<NUM>, to cool to as low of a temperature as possible as it warms the refrigerant flowing counter-currently in the same heat exchanger. The cool refrigerant stream, P-<NUM>, enters an expansion valve or expansion turbine, E-<NUM>, where the pressure drops. The resulting stream vaporizes in E-<NUM> and warms as it cools stream P-<NUM> that comes from the solids-forming heat exchanger. The gaseous refrigerant stream, P-<NUM>, returns through both recuperating heat exchangers, E-<NUM> and E-<NUM>, as it warms back to near ambient temperatures, at which point it completes the loop.

The primary purpose of this loop is to shift the heat of melting CO<NUM> from its nominal melting point, -<NUM>, to a temperature low enough to provide cooling for the solids-forming heat exchanger. This could be enough energy to freeze all of the CO<NUM> in the solids-forming heat exchanger. However, if process conditions and heat losses require additional cooling, a supplementary cooling loop of similar design but without the melting heat exchanger and using either a reverse Brayton or reverse Rankine cycle with recuperative heat recover would supplement this loop. Such cooling loops are well known to one skilled in the art of such processes.

The low-temperature-heat-exchange loop supplements any auto-refrigerant cooling done by dropping the pressure in V-<NUM> and uses a nearly closed-loop refrigeration system to form solids in direct-contact heat exchanger, E-<NUM>. It involves the major pieces of equipment identified in Table <NUM> and the associated connecting streams and controls.

A light gas minimally soluble in the liquid natural gas (the "process stream") in the solids-forming heat exchanger, E-<NUM>, cools in the low-temperature heat exchanger, E-<NUM>, to as low of a temperature as P-<NUM> reasonably allows. This cold gas passes by the liquid natural gas either as a bubbling stream or in a cryogenic spray tower configuration in E-<NUM>. The gas warms as it cools the liquid stream coming from P-<NUM>, solidifying CO<NUM> as the liquid/slurry temperature drops. A fan or other suitable pressurizer provides the pressure rise required for the gas with small amounts of natural gas vapor to return to the low-temperature heat exchanger, E-<NUM>. Gases of potential interest include N<NUM>, any noble gas, CO, and any other suitable material. A small amount of this gas will inevitably be entrained by or dissolve in the stream flowing from the solids-forming heat exchanger, P-<NUM>. This will have to be made up over time by some gas supply. Air is a potential candidate for the gas, as the stream P-<NUM> and all subsequent natural-gas-containing streams should be well above the higher flammability limit and P-<NUM> and all other cooling streams should be well below the lower flammability limit. The circulating gas in P-<NUM>, P-<NUM>, and P-<NUM> should saturate with both methane and CO<NUM> and will neither remove nor add either material from the process, under steady-state conditions, once the initial transient is over.

The following description of a system and method conducted at low pressure is provided for exemplary purposes only and does not fall within the scope of the appended claims.

A similar exemplary process to that illustrated in <FIG> and <FIG> and that operates at pressures too low or temperatures too high to condense the primary component (natural gas, syngas (CO and H<NUM>), or any other light gas mixture in which the major product does not condense at the temperature and pressure of operation) appears in <FIG>. It is similar to the condensing process described above. The melting loop is essentially identical. The major components of the process appear in Table <NUM>. This process is described in terms of its subprocesses similar to above. As with the high-pressure, condensing system, this process begins by conditioning the gas (the "process stream" of this example) to as low of a temperature as is locally achievable using local cooling water or air and condensing heat exchangers (the "first temperature" and the "first pressure" of this example). Any moisture in the process stream is also removed using absorbents or other available equipment.

If the level of CO<NUM> solids formation requires more cooling than is available from this stream, a supplemental refrigerant loop could be incorporated into this low-temperature loop.

The process (not fallin within the scope of the appended claims) begins with raw gas (the "process stream" at the "first pressure" of this example). This gas cools to ambient or slightly below ambient temperature (the "first temperature" of this example) using cooling water or other resources and is then dried using conventional techniques that are not illustrated. The raw, dry gas feed stream first cools in the first heat exchanger E-<NUM> to temperatures near the melting point of pure CO<NUM> at the pressure of the system (the "second temperature" of this example), which in most cases will be near -<NUM>. The first heat exchanger condenses vapors such as natural gas liquids, if present, as the stream cools. These liquids separate (the "first separated liquid stream" of this example) from the gaseous bulk of the flow in E-<NUM> and combine with other liquids streams originating from other portions of the process to form stream P-<NUM>, which returns counter-currently through the heat exchanger to return to near the stream initial temperature as it helps cool the incoming process stream. A liquid pump, included as part of E-<NUM>, raises the pressure of the stream P-<NUM> such that it can mix with stream P-<NUM> to form stream P-<NUM>. The liquid stream may be at sufficient pressure that none of it, or only the lightest components of it, will vaporize as it warms back to its initial temperature. In other examples, some or all of the stream may vaporize.

The bulk of the gas stream (the "process stream") flows as a gas from E-<NUM> into a second heat exchanger, E-<NUM>, that drops the temperature to above and near the frost point (the "third temperature" of this example, i.e., the temperature at which solid CO<NUM> begins to form in the stream). Some trace impurities, such as Hg, will largely condense with the liquids in E-<NUM> while others, such as H<NUM>S, may not condense until later in the process, depending on their concentrations in the raw gas. Some constituents, such as H<NUM>, NO, and CO, will not condense at any stage of this process and are considered light gaseous products produced with the processed gaseous products. Additional vapors may condense as the stream cools, in which case stream P-<NUM> enters the desublimating heat exchanger, E-<NUM>, as a two-phase (gas, liquid) system.

The desublimating heat exchanger may be one disclosed in <CIT> or <CIT>. The desublimation heat exchanger may be a direct-contact heat exchanger and configured to form solids in the heat exchanger without fouling or plugging and may integrate heat transfer to minimize both pressure drop and energy consumption. The gases in stream P-<NUM> (the "process stream") flow counter-currently with colder liquids introduced into the vessel as stream P-<NUM>, and solid CO<NUM> forms on the liquid surface as the streams pass, removing CO<NUM> from the gaseous stream. The slurry that results flows through P-<NUM> into a solid-liquid separator, E-<NUM>. The separator forms a high-solids-loading slurry with the solids which exits in P-<NUM>, and a liquid phase (the "purified contact liquid stream" of this example) which exits in P-<NUM>. This liquid phase includes the liquids introduced into the desublimating heat exchanger in P-<NUM> and those that condensed and entered through P-<NUM>. A portion of this liquid stream splits off as P-<NUM> and the remainder flows back through the heat exchangers as P-<NUM> (the "separated contact liquid stream" of this example). The split portion P-<NUM> passes through a valve or turbine, V-<NUM>, to decrease the pressure and reduce the tendency of solids to form as it cools in a heat exchanger (part of the "first refrigeration system" of this example) that drops the slip stream to the lowest temperature of the system, E-<NUM>. This stream returns to the desublimating heat exchanger, E-<NUM>, and completes the loop from which it started. Gases (the "process stream") that contain no condensable vapors, such as many syngases and flue gases, return all of the liquids (the "contact liquid") back to the desublimating heat exchanger, E-<NUM>.

As an alternative to the processes involved in P-<NUM> and V-<NUM>, the low-temperature loop can use any traditional external refrigerant cooling system if the CO<NUM> content of the condensed phase is low enough. <FIG> illustrates one method to lower the dissolved CO<NUM> content. The CO<NUM> slurry from the desublimating heat exchanger (bubbler or spray tower, P-<NUM>) passes through a solid-liquid separator as usual (E-<NUM>), producing one stream comprising primarily CO<NUM> solid and a second liquid stream. The liquid stream comprises primarily the contacting liquid with some dissolved CO<NUM>. The liquid stream enters a low-pressure gas separation unit (part of the "desaturation system" of this example) in which the pressure is low enough that some dissolved CO<NUM> changes phase. This decreases the dissolved CO<NUM> content and changes the liquid temperature. Some of the dissolved CO<NUM> and the remaining liquid are now in separate phases and return to their original pressures, such as via pumps, in separate streams. The CO<NUM> stream will commonly form a solid in this process, and this stream combines with the previously formed solids stream. The liquid stream that is no longer saturated in CO<NUM> (the "further purified contact liquid stream" of this example) enters a traditional heat exchanger where it cools without risk of solid CO<NUM> formation.

The cooling loop compresses the refrigerant in compressor E-<NUM> and cools it in a heat exchanger, E-<NUM>, to or near the available heat rejection temperature of the process, which will usually be near ambient temperature. Many embodiments will use staged compressors and heat exchangers to maximize the efficiency of this process, effectively combining E-<NUM> and E-<NUM> in several stages in a single unit. The pressurized refrigerant in P-<NUM> cools to near the melting point of CO<NUM> in heat exchanger E-<NUM> and then flows through P-<NUM> and condenses in the melting heat exchanger, E-<NUM>, as it melts the solid CO<NUM> in the gas treatment loop. The resulting liquid refrigerant stream, P-<NUM>, flows through a second recuperating heat exchanger, E-<NUM>, to cool to as low of a temperature as possible as it warms the refrigerant flowing counter-currently in the same heat exchanger. The cool refrigerant stream, P-<NUM>, enters an expansion valve or expansion turbine, E-<NUM>, where the pressure drops. The resulting stream vaporizes in E-<NUM> and warms as it cools the liquid stream P-<NUM> (the "purified contact liquid stream") that comes from the solids-forming heat exchanger. The gaseous refrigerant stream, P-<NUM>, returns through both recuperating heat exchangers, E-<NUM> and E-<NUM>, as it warms back to near ambient temperatures, at which point it completes the loop.

Any of the processes disclosed herein may require additional cooling loops as are common in this industry to overcome heat losses, to balance heat loads, or to otherwise maintain efficient and effective temperature profiles throughout the system.

One of ordinary skill in the art, having the benefit of this disclosure, would understand that a number of variations could be made to the processes of <FIG>. For example, portions of the above processes can retrofit existing natural gas and syngas systems without building the entire process illustrated in the figures. Specifically, using the desublimation heat exchangers and the surrounding equipment can retrofit existing natural gas processes, such as amine processes. In another example, when the primary component is natural gas, a portion of the condensed natural gas formed in some of these processes can be used as liquefied natural gas (LNG) product, which will reduce the amount of material circulating back through the system and change the amount and location of supplemental refrigeration required. Subprocess implementations could be among the first implementations of the embodiments disclosed herein in commercial practice.

In regard to the temperature calculations discussed herein, the high-pressure methods and systems pertain to conditions at which the primary component of the gas condenses at the pressure of the system. Since the gas process streams generally are mixtures, this condensation occurs over a temperature range rather than at a specific temperature as occurs for pure components. Standard theoretical techniques or experimental measurements may be used to determine these temperatures. The predictive techniques in general may include non-ideal equations of state (Peng-Robinson, Soave-Redlich-Kwong, and Predictive Soave-Redlich-Kwong are common examples). Additionally, activity coefficient models (NRTL, Wilson, Uniquac, and Unifac are common examples) may also be used and possibly a Poynting correction. Many of these equations require or would benefit from experimentally determined interaction parameters. In general, the process removes moisture, cools to the point of the major phase beginning to condense, during which cooling minor components such as natural gas liquids substantially condense, cool further as the major component condenses, and cools further in a desublimation heat exchanger as CO2 condenses and usually at least partially forms a solid. These condensation points can all be measured experimentally or predicted as previously indicated.

In the low-pressure process (not falling within the scope of the appended claims), the gases cool to near the temperature at which CO<NUM> begins to desublimate. During this cooling, minor components often condense from the gas. These components can be either separated from the stream and returned, as illustrated, or combined with the separated stream if this is beneficial. Further cooling forms condensed CO<NUM>. Desublimation heat exchangers provide the mechanism for cooling as solid CO<NUM> (or any other component) forms. These condensation points can all be measured experimentally or predicted as previously indicated.

The foregoing cryogenic purifying methods and systems may be used to remove carbon dioxide from raw natural gas that has been dried or comes sufficiently dry from the wellhead. Removal of acid gases, such as carbon dioxide, is often an early stage in natural gas processing. Extractive distillation methods and systems are also disclosed herein that may be used in later stages of natural gas processing. Conventionally, one of the final natural gas processing steps is the separation of natural gas liquids (NGL) from the gas. This is generally done with a demethanizer distillation column that separates methane gas from the NGL. Possible NGL that may be present in raw natural gas include ethane, propane, butane, pentane, and heavier hydrocarbons. Conventionally, the NGL are separated into individual components via additional distillation columns, and each component is further purified. All natural gas constituents absorb CO<NUM> to some degree when in the liquid phase. The existence of a minimum-temperature azeotrope between ethane and carbon dioxide particularly complicates CO<NUM> separation from ethane.

<FIG> (not falling within the scope of the present claims) illustrates a direct sequence of extractive distillation columns (a variation of the conventional scheme) used to separate CO<NUM> from ethane, and is based on the information provided by Luyben and Tavan et al. [<NUM>],[<NUM>]. <FIG> displays the temperature and liquid composition profiles for this base case. CO<NUM> collects as the top distillate from the first distillation column (extraction column), while the bottom product - consisting of ethane and heavier hydrocarbons (C<NUM>+) and substantially free of CO<NUM> - feeds the second distillation column (solvent recovery column). High-purity ethane is obtained as the distillate product, and heavier hydrocarbons (NGL) are obtained as the bottom product of the solvent recovery column. The recovered NGL is divided into two parts, one of which is pumped back into the first column for breaking the azeotrope, and the second part goes to a sequence of conventional distillation columns (not shown) for separation into C<NUM>, iC<NUM>, nC<NUM>, iCs, and nCs product streams.

"Substantially free" as used herein means free of less than <NUM>% of the stated compound, but can include less than <NUM>% and less than <NUM>% of the stated compound on a mole percent basis.

In contrast to the scheme of <FIG>, in some examples of methods of separating carbon dioxide from ethane not falling within the scope of the appended claims, the methods comprise providing a feed stream comprising carbon dioxide, ethane, and higher molecular weight hydrocarbons and introducing the feed stream into a first distillation column, wherein the first distillation column is sized and configured to provide a first distillate stream and a first bottoms stream, wherein the first distillate stream comprises substantially pure carbon dioxide. The methods may further comprise introducing at least a portion of the first bottoms stream into a second distillation column, wherein the second distillation column is sized and configured to provide a second distillate stream and a second bottoms stream, wherein the second bottoms stream comprises higher molecular weight hydrocarbons and is substantially free of carbon dioxide and ethane. The methods may further comprise introducing at least a portion of the second bottoms stream into the first distillation column as a solvent stream, separate from the feed stream. The methods may further comprise introducing at least a portion of the second distillate stream into a third distillation column, wherein the third distillation column is sized and configured to provide a third distillate stream and a third bottoms stream, wherein the third bottoms stream comprises ethane and is substantially free of carbon dioxide and higher molecular weight hydrocarbons. The methods may further comprise combining at least a portion of the third distillate stream into the first distillation column separate from the feed stream.

The higher molecular weight hydrocarbons may include propane, butane, pentane, any natural gas component less volatile than ethane, or combinations thereof.

The first distillate stream may comprise at least about <NUM>% pure, about <NUM>% pure, about <NUM>% pure, about <NUM>% pure, about <NUM>% pure, or about <NUM>% pure carbon dioxide. Additionally or alternatively, the first distillate stream may be condensed and at least a portion refluxed to the first distillation column. The carbon dioxide may be stored for reinjection into the ground or used in other ways.

The first bottoms stream may comprise a significant amount of carbon dioxide impurity, for example, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, or about <NUM> wt% or more. This is in contrast to conventional extractive distillation schemes, wherein the bottom stream of the first distillation column is substantially free of CO<NUM>, such as in the <FIG> scheme where the CO<NUM> is about <NUM> wt% in that stream. Under conventional schemes, the first distillation column is sized and configured to achieve complete CO<NUM> separation in the first column. It has been discovered that by adding a third distillation column and instead of focusing on complete CO<NUM> separation, recycling an azeotropic mixture of CO<NUM> and ethane from the third distillation column, surprisingly, capital costs and operating costs can be reduced, as compared to the two-column scheme of <FIG>.

In keeping with the foregoing, the second distillate stream may comprise ethane and a significant amount of carbon dioxide impurity, such as, for example, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, about <NUM> wt% or more, or about <NUM> wt% or more. The second distillate stream may be partially condensed and refluxed to the second distillation column and the uncondensed portion introduced to the third distillation column.

The third distillate stream will generally comprise a substantial amount of ethane and a substantial amount of carbon dioxide, such as an azeotropic mixture of the two. The third distillate stream may be partially condensed and refluxed to the third distillation column and the uncondensed portion is the portion combined with the feed stream.

As should be understood, the feed stream may comprise the bottoms stream from a demethanizer column (not illustrated). One of ordinary skill in the art, with the benefit of this disclosure, would understand how to optimize the ratio of the second bottoms stream (solvent stream) to the feed stream so as to optimize the process. The portion of the second bottoms stream not combined with the feed stream may be further processed using conventional processes. For example, the portion of the second bottoms stream not used as solvent may be sent to a depropanizer column for separating out propane. The bottoms from the depropanizer column may be sent to a debutanizer column for separating out butane and so forth for separating out pentane.

In some embodiments, heat may be exchanged between the second bottoms stream and the third bottoms stream and/or the third distillate stream prior to introducing a portion of the second bottoms stream as a solvent.

In some embodiments, such as in the illustrated embodiments, heat is not exchanged between the first distillate stream and the second bottoms stream, thereby allowing the CO<NUM> in the first distillate stream to remain in the liquid phase. The methods of separating carbon dioxide from ethane contemplated herein may also be used with a new installation or with an existing extractive distillation process that has been retrofitted. Accordingly, methods of retrofitting are contemplated. In some examples, methods of retrofitting (not falling within the scope of the appended claims) an existing extractive distillation process, such as that disclosed in <FIG>, comprise: reconfiguring the first distillation column to provide a first bottoms stream comprising a significant amount of carbon dioxide impurity; introducing at least a portion of a second bottoms stream of the second distillation column as a solvent stream to the first distillation column, separate from the feed stream; and adding a third distillation column and introducing at least a portion of a second distillate stream of the second distillation column into the third distillation column, wherein the third distillation column is sized and configured to provide a third distillate stream and a third bottoms stream, wherein the third bottoms stream comprises ethane and is substantially free of carbon dioxide and higher molecular weight hydrocarbons.

The retrofit methods (not falling within the scope of the appended claims) may further comprise combining at least a portion of the third distillate stream with the feed stream prior to introducing it to the first distillation column. The third distillate stream may be partially condensed and refluxed to the third distillation column and the uncondensed portion combined with the feed stream. The third distillate stream may comprise an azeotropic mixture of ethane and carbon dioxide.

In some examples (not falling within the scope of the appended claims), an extractive distillation system for separating ethane from carbon dioxide comprises a first distillation column configured to receive a feed stream and a separate recycle stream. The first distillation column is sized and configured to provide a first distillate stream and a first bottoms stream. The system further comprises a second distillation column configured to receive the first bottoms stream, wherein the second distillation column is sized and configured to provide a second distillate stream and a second bottoms stream. The system may include a diversion system configured to provide a portion of the second bottoms stream as the recycle stream. The system further includes a third distillation column configured to receive the second distillate stream, wherein the third distillation column is sized and configured to provide a third distillate stream and a third bottoms stream, wherein the third distillate stream is combined with the feed stream.

Depending on the composition of the feed stream, the feed stream may be introduced within the lower <NUM>% of the first distillation column. Likewise, depending on the composition of the feed stream, the recycle stream may be introduced within an upper <NUM>% of the first distillation column and may be substantially condensed. The first bottoms stream may in turn be introduced to the second distillation column within <NUM>% of the middle of the second distillation column. The second distillate stream may be introduced to the third distillation column in an upper portion of the third distillation column and may be substantially uncondensed.

<FIG> illustrates an exemplary system and method (not falling within the scope of the appended claims) for separating carbon dioxide and ethane, as discussed above. The exemplary system involves three columns: the CO<NUM> recovery or extraction column ("first distillation column"), the solvent recovery column ("second distillation column"), and the concentrator column ("third distillation column"). In this process, not all the CO<NUM> exits the top of the extraction column, the first column. In this example, the bottom product of the extraction column contains <NUM> mol % CO<NUM> along with ethane and heavier hydrocarbons. The second column recovers high-purity solvent. The recovery column distillate feeds the concentrator column, which produces ethane as a product and an azeotropic mixture recycle stream.

The exemplary extractive column had <NUM> stages and operated at <NUM> atm. The feed gas entered on tray <NUM> and the solvent with the flow rate of <NUM> kmol/s (Solvent/Feed = <NUM>) entered on tray <NUM> near the top. <FIG> plot the temperature and liquid composition profiles in the extractive column, respectively. The solvent in the extractive distillation column ("first distillation column") altered the relative volatility between CO<NUM> and ethane, driving CO<NUM> to the top of the column and ethane to the bottom of the column. The upper section of the first column (above the entrainer (i.e., solvent) feed location) separated the CO<NUM> and the entrainer. <FIG> indicates that the CO<NUM> concentration increased at the entrainer entry point (stage <NUM>). The middle of the column, between the entrainer feed stage and the fresh feed stage, prevented ethane from going up the column. <FIG> clearly shows that the concentration of ethane increased from stages <NUM> to <NUM>, where the entrainer and the feed enter, respectively. The bottom of the column, below the fresh feed location, prevented CO<NUM> from going down the column. The extractive column produces a CO<NUM>-rich distillate (<NUM> mol %).

<FIG> illustrates the effects of changing the solvent flow rate (S) and/or reflux ratio (RR). Increasing the reflux ratio decreased the impurity of solvent in the distillate, while increasing the solvent flow rate improved the distillate CO<NUM> purity. However, the same is not true for the CO<NUM>-ethane system when using the NGL solvent. The effect of higher solvent flow rates depended on the reflux ratio; in the higher RR range, increasing the solvent flow rate also increased the distillate CO<NUM> purity (<FIG>). However, in the lower RR range, the opposite occurred.

Additionally, <FIG> reveal nonmonotonic relationships between RR and both of the distillate impurities (C<NUM> and C<NUM>), but they were opposite in shape. The C<NUM> curve reaches a minimum and the C<NUM> curve (the lightest of the components in the NGL solvent) reaches a maximum. The same relationships are also true of the solvent flow rate: more solvent decreased C<NUM> impurity but increased C<NUM> impurity.

These interesting phenomena occurred because the solvent is chemically similar to the ethane being separated from the CO<NUM>. While operated at the lower of the two possible reflux ratios, the minimum solvent flow rate was found which met the two specifications for the extractive column. As shown in <FIG>, the solvent flow rate was <NUM> kmol/s and the reflux ratio was <NUM>. The resulting heat exchanger duties were <NUM> MW in the condenser and <NUM> MW in the reboiler. The distillate flow rate was <NUM> kmol/s of mostly CO<NUM> (<NUM> mol %) with impurities of <NUM> mol % ethane and <NUM> mol % propane. The bottoms flow rate of the column was <NUM> kmol/s and carried most of the ethane, some of the CO<NUM> in the fresh feed, and heavier hydrocarbons to the solvent recovery column ("second distillation column").

The distillate of the extractive column in this design remains liquid. In contrast, the distillate of the conventional design cools the recycled NGL, which converts it to a vapor stream. There are several potential heat integration steps for this liquid stream, depending on the potential use of the CO<NUM>. For example, a simple pump could pressurize the stream to above its vapor pressure at ambient conditions. The stream could then warm to ambient temperature by heat integration with any one or several of the condensers, further decreasing the process energy demand. The final liquid stream would then be suitable for enhanced oil recovery or other pipeline-based, large-scale CO<NUM> applications. Alternatively, if the CO<NUM> was to be locally vented, the stream could reduce the overall process energy demand significantly by heat integration with one or more of the condenser circuits.

The solvent recovery column ("second distillation column") had <NUM> stages, and the first bottoms stream of the extractive column (first column) was fed on tray <NUM>. Unlike the conventional design that uses a total condenser, this column in this exemplary design had a partial condenser. The design specifications of this column were <NUM> mol % propane in the distillate and <NUM> mol % ethane in the bottoms. A reflux ratio of <NUM> was used to achieve these specifications.

<FIG> exhibit the temperature and liquid composition profiles in the solvent recovery column, respectively. The ethane and CO<NUM> concentrations functionally monotonically decreased from the distillate to the bottoms, but the C<NUM> profile had two local maxima, one each between the feed and the distillate and the feed and the bottoms, with overall increasing concentration from the distillate to the bottoms.

The condenser duty was <NUM> MW, and the reboiler duty was <NUM> MW. The heavy hydrocarbons exited with the bottoms product ("second bottoms stream") and split, via a diversion system, into the NGL product (<NUM> kmol/s) and the solvent, which recycles back to the extractive column separate from the feed stream. As discussed previously, the NGL stream could pass through a sequence of traditional distillation columns for propane, butane, and pentane recoveries, which were not included in any of these simulations. The distillate of the solvent recovery column was a mixture of CO<NUM> (<NUM> mol %) and ethane (<NUM> mol %) that needed to be concentrated before recycling to the initial feed.

The concentrator column ("third distillation column") had <NUM> stages, and the distillate of the recovery column was fed on tray <NUM>. This column also had a partial condenser. Ethane with high purity (<NUM> mol %) formed the bottoms product (<NUM> kmol/s) after heat recovery. The mixture of CO<NUM>-ethane went overhead with a molar flow rate of <NUM> kmol/s with the CO<NUM> concentrated up to <NUM> mol %. After heat recovery, this was recycled back to the extraction column as part of the feed stream.

To establish a fair comparison between the exemplary <FIG> process (not falling within the scope of the appended claims) and the exemplary <FIG> process (also not falling within the scope of the appended claims), the sequential quadratic programming (SQP) method was used to optimize both designs. The SQP optimization method coupled with the efficient sensitivity analysis tool from Aspen Plus minimized the total energy requirement for the extractive column, as follows: <MAT> where the optimization parameters used here are: the total number of stages (NT), feed location (NF), solvent location (NS), reflux ratio (RR), boilup rate (V), and solvent to feed flow rate ratio ( <MAT>), while ym and xm are the vectors of the obtained and desired purities for the m products, respectively. An additional objective function determined the optimal energy cost vs. the number of stages [<NUM>]-[<NUM>].

The total capital costs, total operating costs, and total annual costs (TAC) were calculated. TAC was calculated as <MAT> where a plant life of <NUM> years was considered in this investigation. This somewhat exceeds many natural gas process plant lifetimes but also allows for equipment reuse. Most of the savings in the <FIG> process were in operating costs, so the economics are relatively insensitive to the plant cost. The Aspen Process Economic Analyzer provides relative costs, taking into account capital and operating costs together with technical and process parameters. Table <NUM> compares certain economic performance indicators of the two processes. The introduction of the third column caused a noticeable decrease in the reflux ratios of the extractive and recovery columns, which was <NUM> and <NUM>, respectively, compared with <NUM> and <NUM> in the <FIG> system. The capital costs of the columns were significantly altered by the reflux ratios, recycle flow rates, and entrainer usage in the distillation cases. Additionally, these parameters determined the process heat duties and product quality. As a result, the <FIG> process leads to <NUM>% lower TAC, with a <NUM>% reduction in specific energy demand and CO<NUM> emissions. All costs indicated here decrease. Total capital costs and most operating costs decrease by about <NUM>%, with a <NUM>% reduction in steam costs, consistent with the sum of the reboiler thermal loads being the biggest change.

The energy requirements closely correlate with CO<NUM> emissions when no heat integration is considered. When part of the process heat is reused instead of primary energy, then the CO<NUM> emissions are lower as compared to the figure expected from the energy data [<NUM>],[<NUM>]. Fuel combustion calculations assumed that air was in excess, which ensures complete combustion and prevents formation of carbon monoxide. The amount of CO<NUM> emitted was related to the amount of fuel burned in the heating devices, and was calculated according to the method described by Gadalla et al. and Kiss et al. [<NUM>],[<NUM>]: <MAT> where α = <NUM> is the ratio of molar masses of CO<NUM> and C, and NHV is the net (lower) heating value of natural gas with a carbon mass fraction of <NUM>. Hence the quantity of fuel used was calculated as follows: <MAT> where λproc (kJ/kg) and hproc (kJ/kg) are the latent heat and enthalpy of the steam, and TFTB (K) and T<NUM> (K) are the flame and stack temperatures, respectively. This equation represents a steam balance around the boiler and relates the quantity of fuel necessary to provide a heat duty of Qproc' The hourly rate of CO<NUM> emissions for the <FIG> process and the <FIG> process also appear in Table <NUM>. This indicates that the exemplary process and system decreases the carbon footprint for extractive distillation of ethane and carbon dioxide.

As compared to the <FIG> process (not falling within the scope of the appended claims), the exemplary process strategy (also not falling within the scope of the appended claims) showed an approximately <NUM>% reduction in total energy demand and associated carbon emissions, most of which were from reduced steam demand. Aspen Plus process economics analyses indicated about a <NUM>% reduction in capital and a <NUM>% reduction in operating costs when comparing optimized versions of the <FIG> process and the <FIG> process. The <FIG> process reduced the total annual costs (TAC) by <NUM>%, without compromising the desired purification. The <FIG> process was also easier to operate because it was unnecessary to withdraw CO<NUM> completely in the extractive column. Additionally, the <FIG> process produced CO<NUM> as a liquid product, which avoided the significant amount of energy required for liquefaction.

As discussed above, methods of cryogenically purifying process streams may be used to remove CO<NUM> from raw natural gas. Likewise, methods of separating carbon dioxide from ethane may be used in purifying the ethane separated from natural gas. Accordingly, methods are contemplated herein that encompass both processes. A method of separating carbon dioxide from a process stream may comprise providing a process stream comprising gaseous carbon dioxide and a primary component. The process stream may further comprise at least a secondary component. The method comprises cooling the process stream to at or below the frost point of carbon dioxide in the process stream, thereby forming solid carbon dioxide and then physically separating solid carbon dioxide from the process stream to form a separated solid carbon dioxide slurry stream and a liquid primary component stream. The method may further comprise separating at least the secondary component and residual carbon dioxide from the liquid primary component stream. The method may further comprise separating the residual carbon dioxide from the secondary component using extractive distillation with a concentrator distillation column. In the natural gas process setting, the primary component may be methane and the secondary component may be ethane. This example of combined processes may include any of the features of the separate cryogenic and extractive distillation processes disclosed in more detail above.

It should be understood that just as the cryogenic processes may be applied to non-natural gas process streams, likewise, the extractive distillation processes may be applied to streams comprising ethane and CO<NUM> that originate from sources other than natural gas. For example, if a process stream comprising ethane and CO<NUM> did not also contain higher molecular weight hydrocarbons that could be used as a solvent, the extractive distillation process could still be used. A solvent such as butane could be introduced and completely recycled through the process, instead of using naturally present higher molecular weight hydrocarbons.

Claim 1:
A method of separating carbon dioxide from a process stream comprising:
providing a process stream comprising carbon dioxide and a primary component;
wherein the process stream is gaseous carbon dioxide;
wherein the process stream is at a first temperature and a first pressure; and
cooling the process stream to at or below the condensation temperature of the primary component in the process stream;
separating any gases from the process stream that did not condense during cooling the process stream to at or below the condensation temperature of the primary component to form a first separated gaseous stream;
cooling the process stream further to at or below the frost point of carbon dioxide in the process stream, thereby forming solid carbon dioxide; and
separating physically solid carbon dioxide from the process stream to form a separated solid carbon dioxide slurry stream and a liquid primary component stream;
characterized in that cooling the process stream to at or below the frost point of carbon dioxide in the process stream comprises:
directly contacting the process stream with a colder contact gas stream, and
separating a warmed contact gas stream from a slurry of the condensed primary component and the solid carbon dioxide, and
wherein the contact gas comprises nitrogen, air, a noble gas, carbon monoxide or a combination thereof.