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5.4 Treatment plants for gas production

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VOLUME I / EXPLORATION, PRODUCTION AND TRANSPORT 5.4.1 Introduction This chapter examines the treatments required for the transport and distribution of natural gas. It will discuss the main processes used to purify gas and the requisite equipments both for gas bearing produced by gas reservoirs, and gas associated with crude oil. In the specific case of production from gas condensate reservoirs, the vapour phase and its treatment will be analysed in this chapter, whereas the liquid phase and its treatment have already been discussed in Chapter 5.3. This discussion covers surface plants comprised between the production wellheads (excluded) up to and including delivery to the gas pipelines used for transport and distribution. These treatment plants include gas/liquid separators, dehydration, condensate removal and sweetening units, and those used to remove other compounds in the gas such as mercaptans and mercury. Considering the significant increase in the transport of liquefied natural gas, this chapter will describe the pretreatments needed for liquefaction, in addition to liquefaction processes. The chemical and physical properties of natural gas depend on its origins and composition. The latter, except in unusual cases, does not significantly influence either treatment or transport. For this reason, the following discussion will not take into consideration the origins of the gas. Specifically, the chemical composition and the greater or lesser percentages of heavier hydrocarbons present in the gas have no impact on the considerations relating to the processes under examination. Therefore no distinction will be made between dry gas, gas condensate and gas associated with crude oil. However, it is important to note that the surface facilities needed to handle and purify natural gas are not always identical to those used to treat refinery gases and synthetic gases. Natural gas contains higher hydrocarbons, ranging from paraffins to aromatics and naphthenes; however, in contrast to refinery gases and synthetic gases, it does not contain olefins. As far as contaminants are concerned, natural gas may contain a wide range of compounds which can confer upon it negative properties, both in terms of transport and distribution. The main non-hydrocarbon component of natural gas is water. As will be seen in the following discussion, this is removed by dehydration, the most common treatment process in natural gas production. Chapter 5.3 dealing with oil treatment examined the properties of the water present in a reservoir fluid. In the specific case of gas, formation water has lower salinity than that commonly found in oil fields; despite this, the presence of chlorides in the water produced alongside gas must be carefully evaluated and treated to reduce to a minimum or avoid completely the contamination it causes during the purification processes undergone by the gas. Whereas the presence of modest quantities of free water in an oil is acceptable, in gas this component must be completely removed to avoid the formation of condensation under the most critical transport and distribution conditions, in other words at high pressure and low temperature. The parameter defining this condition is the water dew point, that is the temperature, at a given pressure, at which the first drop of condensed water forms. Natural gas may also contain inert gases such as nitrogen, N 2 and helium, He. In general their presence is acceptable within the limits of variability of the calorific value. Natural gas often also contains carbon dioxide, CO 2 , and this compound is tolerated at higher or lower 681 5.4 Treatment plants for gas production
Transcript

VOLUME I / EXPLORATION, PRODUCTION AND TRANSPORT

5.4.1 Introduction

This chapter examines the treatments required for thetransport and distribution of natural gas. It willdiscuss the main processes used to purify gas and therequisite equipments both for gas bearing producedby gas reservoirs, and gas associated with crude oil.In the specific case of production from gascondensate reservoirs, the vapour phase and itstreatment will be analysed in this chapter, whereas theliquid phase and its treatment have already beendiscussed in Chapter 5.3.

This discussion covers surface plants comprisedbetween the production wellheads (excluded) up toand including delivery to the gas pipelines used fortransport and distribution. These treatment plantsinclude gas/liquid separators, dehydration, condensateremoval and sweetening units, and those used toremove other compounds in the gas such asmercaptans and mercury.

Considering the significant increase in thetransport of liquefied natural gas, this chapter willdescribe the pretreatments needed for liquefaction, inaddition to liquefaction processes.

The chemical and physical properties of naturalgas depend on its origins and composition. The latter,except in unusual cases, does not significantlyinfluence either treatment or transport. For this reason,the following discussion will not take intoconsideration the origins of the gas. Specifically, thechemical composition and the greater or lesserpercentages of heavier hydrocarbons present in the gashave no impact on the considerations relating to theprocesses under examination. Therefore no distinctionwill be made between dry gas, gas condensate and gasassociated with crude oil.

However, it is important to note that the surfacefacilities needed to handle and purify natural gas are

not always identical to those used to treat refinerygases and synthetic gases.

Natural gas contains higher hydrocarbons, rangingfrom paraffins to aromatics and naphthenes; however,in contrast to refinery gases and synthetic gases, itdoes not contain olefins.

As far as contaminants are concerned, natural gasmay contain a wide range of compounds which canconfer upon it negative properties, both in terms oftransport and distribution.

The main non-hydrocarbon component of naturalgas is water. As will be seen in the followingdiscussion, this is removed by dehydration, the mostcommon treatment process in natural gas production.

Chapter 5.3 dealing with oil treatment examinedthe properties of the water present in a reservoir fluid.In the specific case of gas, formation water has lowersalinity than that commonly found in oil fields; despitethis, the presence of chlorides in the water producedalongside gas must be carefully evaluated and treatedto reduce to a minimum or avoid completely thecontamination it causes during the purificationprocesses undergone by the gas.

Whereas the presence of modest quantities of freewater in an oil is acceptable, in gas this componentmust be completely removed to avoid the formation ofcondensation under the most critical transport anddistribution conditions, in other words at high pressureand low temperature. The parameter defining thiscondition is the water dew point, that is thetemperature, at a given pressure, at which the firstdrop of condensed water forms.

Natural gas may also contain inert gases such asnitrogen, N2 and helium, He. In general their presenceis acceptable within the limits of variability of thecalorific value.

Natural gas often also contains carbon dioxide,CO2, and this compound is tolerated at higher or lower

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5.4

Treatment plantsfor gas production

values, subject to the dehydration of the gas itself; thiscompound may confer considerable acidity on the gas,combining with any condensation water present.

By contrast, the presence of hydrogen sulphide,H2S, often present in gas alongside other sulphurcompounds such as mercaptans, is not tolerated.Whereas the contents of the former compound must bereduced to negligible values due to its toxicity, thelimit content of other sulphur compounds (totalsulphur) is higher.

Another substance frequently found in gas ismercury, Hg, which may be present as elementarymercury, or in the form of some of its compounds, ofwhich the most frequent is mercury sulphide HgS. Thepresence of this heavy metal in the gas may causesignificant problems, not so much during productionand in the gas gathering networks, as during subsequentcondensate removal and liquefaction treatments. This isbecause at low temperatures elementary mercurychanges from the gas phase in which it is found in thereservoir to the liquid phase; this condition must beavoided by treating the gas appropriately.

Finally, gas production may contain solid particlesof varying origins, such as sand or colloidal clay,dragged from the reservoir from which it is produced,or formed by corrosion. These suspended solids mustbe removed before the gas can be commercialized.

5.4.2 Sales gas specifications

Calorific value and Wobbe indexMost gas production is used as fuel for civic

purposes, and as such is distributed to users throughappropriate networks.

For this purpose, the calorific value of the gas musthave a limited range of variability. Generally speaking,in analysing the compatibility and interchangeability ofvarious types of gas, the reference parameter used is thegross calorific value of methane (9,001.6 kcal/Sm3),since the latter is by far the predominant component ofnatural gas production. The range of variability of thecalorific value in general must not exceed �10%. Thisfairly approximate factor on its own serves to define theinert gas content (N2, CO2, He, etc.), acceptable for thecommercialization of natural gas.

Natural gas may also contain higher hydrocarbons(ethane, propane, butanes, etc.) characterized by a farhigher calorific value than that of methane; the inertgas and higher hydrocarbon content may therefore leadto considerable variability, whilst still meeting therestriction described above.

In controlling combustion, alongside calorificvalue, another parameter of the gas is also important;its specific density with respect to air. For the sake of

simplicity, it can be stated that the combined effect ofthe two parameters described above on gas combustioncan be expressed by a single parameter known as theWobbe index, and defined as HHV/sg0.5, where HHVis the high heating value and sg the specific gravity ofthe gas mixture under examination. In the case of themethane used as a reference point, the Wobbe index is12,094.8 kcal/Sm3.

In evaluating the compatibility of a gas withdelivery to a distribution network it is more correct torefer to the Wobbe index as defined above, rather thanto the simpler calorific value. The range of variabilitygenerally accepted in the latter case is in the order of�5%. As can easily be seen, this specification isconsiderably more restrictive than the former,particularly due to the resulting limitation determinedby the maximum allowed percentage of inertcomponents.

In treating gas, the specification described above isfundamentally important, since it is the basis for thetreatments which the gas must undergo in order for itto be commercialized.

Water and hydrocarbon dew pointAs mentioned above, gas must be transported

through long gas pipelines and then delivered tousers through distribution networks. The gas istreated upstream the transport and distributionsystem. The distribution network does not includeany treatment, but merely reductions in pressure; as aresult, the gas must meet specifications such as toensure that no conditions under which water and/orhydrocarbons might condense arise during any of thephases of transport and distribution. Obviously, thesespecifications vary depending on the area crossed.Often, water and hydrocarbon dew point conditionsalso depend on the specific subsequent treatmentwhich the gas must undergo. As will be seen below,the gas may be treated to recover higherhydrocarbons, or even liquefied for transport atatmospheric pressure with LNG carriers. In thesecases, the gas must undergo a more or less severecooling process; as a consequence, the requisitereduction of the water and hydrocarbon dew pointmay be more or less severe.

Usually these two specifications consider twodifferent values, although obviously they could beexpressed as a single value. There are two reasons forthis difference. As will be shown below, even verymodest water condensation in a transport pipeline orduring the treatment of gas may cause far moresignificant problems, such as the obstruction of thepipeline itself, than those caused by the modestcondensation of hydrocarbons alone. For this reason, theoperating margins for water condensation must be

682 ENCYCLOPAEDIA OF HYDROCARBONS

DEVELOPMENT PHASE OF HYDROCARBON FIELDS

higher. A second reason is linked to the particularbehaviour of the two-phase gas-liquid hydrocarbonequilibrium, as compared with the gas-waterequilibrium. The latter shows a univocal pattern forsaturation as a function of pressure. Given an identicalwater content, the higher the pressure the higher the dewpoint. In the case of hydrocarbons, this behaviour ismore complex, and saturation and/or condensation donot have univocal behaviour versus pressure. The phasecurve shown in Fig. 1 demonstrates that at low pressuresthe condensation temperature increases as pressureincreases, up to a maximum value (cricondentherm);above this value the variation of the dew point as afunction of pressure is inverted, and an increase inpressure leads to a reduction in the dew point. The rangeof pressures above the cricondentherm is thereforeknown as the retrograde condensation zone (see Chapter4.2). Under this condition, a reduction of pressure atconstant temperature leads to the condensation of liquidhydrocarbons rather than to undersaturation.

The above description explains why, in defininghydrocarbon dew point specifications, reference is notmade to a predetermined pressure, as for water, but tothe full range of pressures, starting from atmosphericpressure. The water dew point specification is morecommonly defined as the water content in the gas.Obviously, the hydrocarbon dew point cannot beexpressed as a predetermined content.

Inert gas content (CO2 and N2)The carbon dioxide (CO2) concentration may be

relatively high without interfering with the mainrestriction relating to the Wobbe index.

Basically, the maximum permitted CO2 content islinked to the maximum total permitted inert gas content(usually CO2�N2). As will be seen in the discussion ofpurification treatments required to obtain the requisite

quality, it is easier to remove carbon dioxide thannitrogen. It is therefore preferred, as far as possible, toleave the latter unaltered and to remove CO2 until therequired Wobbe index and/or maximum total inert gascontent is reached. In general, CO2 contents of up to 2-3% mol are acceptable; the total inert gas content,CO2�N2, should not exceed 6.5% mol.

Maximum hydrogen sulphide (H2S) contentThe maximum permitted content for distribution

networks in the United States and North America is1/4 grain/102 Sft3�5.74 mg/Sm3�4.0 ppm in volume;the same values are generally accepted in Europe. Thisextremely low content means that hydrogen sulphidemust be removed almost entirely.

It is worth noting that gas sweetening processeslead simultaneously to a significant removal of carbondioxide.

Total sulphur contentSince it is not linked to safety problems, but merely

to the reduction of air pollution, the total sulphurcontent, which often coincides with that of mercaptansulphur, is 17 mg/Sm3, and thus far higher than that ofhydrogen sulphide. Many processes for the removal ofhydrogen sulphide and acid gases in general have only amodest impact on the removal of mercaptans, and thusof total sulphur. Very frequently, a natural gas with ahigh hydrogen sulphide content also contains significantquantities of mercaptans, which must be reduced to theaforementioned specification with appropriate treatment.

The same is true for Liquefied Petroleum Gases(LPG) obtained from gas with condensate removaltreatment. As far as hydrogen sulphide is concerned,LPGs must meet a specification similar to that for gas;as far as mercaptans are concerned, by contrast, LPGs,which are also used as vehicle fuels, must meet thesame specifications as gasoline and pass the so-calledDoctor test; in other words they must contain less than10 ppm of mercaptans. As such, in order to becommercialized, LPGs must undergo suitabletreatment, very similar to that described for oil.

Suspended solids contentThere is no uniform specification for solids contents.

The standard for the removal of suspended solids in gasis generally met as a consequence of the removal ofliquid particles. Only in unusual cases, dictated mainlyby specific treatment needs, is the gas subjected tofiltration in order to remove suspended solids.

Mercury contentMany natural gases contain significant amounts of

mercury. As a general rule, this contaminant should beremoved to reduce air pollution; in practice, natural

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pres

sure

cricondentherm

temperature

100806040200

dewpoints

bubblepoints

retrogradecondensation region

C

Fig. 1. Phase diagram for a multi-component mixture.

gas often undergoes only very simple treatmentprocesses such as dehydration. In these cases, alimited mercury content has no significant impact ontreatment itself, and is therefore accepted. This is nottrue when the gas must undergo cooling processes; inthis case a stringent mercury removal treatment, tovalues of 10 ppb, is needed.

5.4.3 Gas/liquid separation

From a theoretical point of view (dynamics ofparticles dispersed in a fluid), there is no differencebetween the setting of drops of liquid in a gas, andliquid/liquid gravity separation. Even from a practicalperspective, the differences are insignificant.

The final settling velocity of the particles, Vt, iscalculated by balancing the forces acting on theparticles during their setting in the dispersing fluid. Inthe case of non-deformable spherical particles ofdiameter Dp, we obtain:

Vt�[4gDp(rp�r) /3C�r]1/2

where g is the acceleration of gravity, rp the density ofthe particles, r the density of the dispersing fluid, andC� the drag coefficient. In the case of the separation ofliquid particles suspended in a dispersing liquid, thesettling of the particles usually occurs in conditions oflaminar flow, given the low value of the Reynoldsnumber (Re), defined by the equation: Re�DpVtr/m(where m is the viscosity of the dispersing fluid). Thecoefficient C� can thus be expressed by the simpleequation: C��24/Re, making it possible to calculate Vtusing Stokes’ law (see Chapter 5.3). By contrast, whendealing with the separation of liquid particles in a gas,Re takes on a far higher value due to the low viscosityof the dispersing fluid (gas), and it is therefore nolonger possible to assume that the particles settleunder conditions of laminar flow. In this case thecoefficient C� is calculated as a function of Re usingsuitable programmes, or more complex equations.Solving the equation linking Vt to Dp requirescalculation by trial and error, and this methodologyhas been codified in API RP 520.

The most obvious use for this methodology is sizingor evaluating the flare KOD (Knock Out Drum)separator. In this specific instance, for safety reasons, itis impossible to use any type of internals suited to aidingthe coalescence of dispersed drops. As a consequence,the sizing of the vessel is based exclusively on thegravity separation described above. Using thismethodology, the diameter of the particles removed fromthe gas stream is, in the case cited, in the order of 250mm, whereas in other cases it may reach slightly lowervalues; however, this is an extremely approximate

separation. Another example of the use of simple gravityseparation is the slug catcher (see Section 5.4.6), wherethe main purpose of the equipment is to block the slugsof liquid entrained in a gas stream.

In all other cases where it is necessary to ensure agood reduction of entrained liquids alongside a roughgravity separation, a demisting unit must be insertedinto the separator to help the drops to coalesce,facilitating the removal of particles with a diameterabove 10 mm. It is not easy to define the degree ofpurification that can be obtained by these means; thestatistical distribution of the diameters of the drops isin fact difficult to determine.

It is common practice to consider the liquidcontent entrained in a gas downstream a demisting unitsized to remove drops with a diameter above 10 mm asbeing equal to 0.1 Gal/106 Sft3. This practicalsimplification derives from the data from numerousseparation units, rather than from studies of particledynamics and the statistics of drop distribution.

One of the most commonly used demisting unit isthe wire mesh pad: this is inserted perpendicular to thegas flow at a predetermined distance from the outflowand inflow nozzles of the stream itself.

In the separation of liquid from gas, the mostfrequently used typology is the vertical separator,where the gas outflow nozzle is positioned at thecentre of the elliptical upper part of the vessel. In thiscase, the sizing of the separator is mainly based oncalculating the minimum section needed for theinstallation described above.

The equation determining the maximum velocity ofthe gas in the demisting wire mesh pad is as follows:

Vt�K[(rL�rG) /rG]1/2

where rL is the density of the liquid particles and rGthe density of the dispersing gas. The constant K variesas a function of the type of separator required and itsworking pressure. The values recommended by theGPSA (Gas Processors Suppliers Association) as afunction of the various working pressures range from0.36 at atmospheric pressure to 0.21 at 100 bar andabove (when the density is espressed in lb/Sft3).

For some specific types of separation, such as theKOD separators used to protect compressors andexpanders, the values of K are lower, with acoefficient of 0.8.

In a vertical separator, the cross-sectional area offlow thus corresponds to that of the horizontal wiremesh pad installed inside it. For other dimensions,such as height, it is necessary to know the flow rate ofliquid to be removed and determine the hold-up time,which may vary depending on operating conditionsand the chemical and physical properties of the fluidsto be separated.

684 ENCYCLOPAEDIA OF HYDROCARBONS

DEVELOPMENT PHASE OF HYDROCARBON FIELDS

In some specific cases, horizontal separators mayalso be used, with the demisting unit positioned in anidentical way to vertical separators, in other wordswith the gas flowing from bottom to top. When theseparator may receive a two-phase flow withsignificant and discontinuous amounts of liquid, adouble-barrel horizontal separator may also be used.

Other types of demisting units are also commonlyused to separate liquids from gas. The most frequentlyused are cyclones, and the so-called vanes. With theformer, coalescence and the resulting separation arebased on the centrifuge effect obtained in the circularmovement of the gas. The results are extremelyinteresting: it is possible to obtain a higher degree ofseparation, with the almost complete removal of particlesas small as 3 mm, or with even smaller diameters, and/ora reduction of the diameter of the separator into whichthe cyclones are inserted. The disadvantage of thissolution lies in the greater pressure drop resulting from itsuse; however, this disadvantage is often negligible.

The second type of coalescence elements arepackages of vanes, also frequently used in oil-gasseparators (see Chapter 5.3). In this type of separation,coalescence is due to the so-called ‘chicane effect’,very similar to the centrifuging described above.

Both systems make the equipment more compact,and even make it possible to separate out any solidparticles entrained by the gas with the liquid in anextremely efficient way. They are also self-cleaning, inother words they allow the solid particles to be drainedaway automatically in the separated liquid. Thecommon wire mesh pad also removes solid particlestogether with the liquid but, unfortunately, given itsengineering properties, it tends to retain the separatedsolids and collapse. When the gas to be separatedcontains a significant amount of solids, such asparaffin crystals which have separated from the liquid,as in the case of a low temperature separator, the wiremesh pad must be replaced with vanes or cyclones.

The materials most commonly used to buildseparators are LTCS (Low Temperature Carbon Steel)for the outer shell, and AISI (American Iron and SteelInstitute) 304L or 316L for the internals. When the gashas a significant acid gas content (H2S and CO2) andis saturated in water, the shell must also be resistant toacid corrosion. This is usually achieved by plating theinterior of the vessel (3 mm thick) with AISI 316L.

5.4.4 Dehydration and condensateremoval

Three-phase vapour/liquid/solid equilibriumIt is known that a hydrocarbon mixture in the

gaseous state may give rise to the condensation of

water and/or hydrocarbons as operating conditionsvary.

When free water forms in a natural gas stream, asan effect of high pressure and low temperature thismay lead to the formation of compounds, hydrates,which are highly unstable but have the physicalproperties of a solid.

The formation of a triple-point equilibrium(vapour, liquid, solid) depends on operating conditionsand the properties of the gas mixture underexamination.

In the presence of free water, light hydrocarbonsfrom methane to isobutane may give rise to theformation of solid hydrates at high pressures and lowtemperatures.

The tendency of light hydrocarbons to form hydratesincreases with the molecular weight of the hydrocarbonup to isobutane; from normal butane onwards thistendency disappears and the component behaves like aninert in the equilibrium described above.

The presence of carbon dioxide and hydrogensulphide may also facilitate hydrates to form; for theformer this tendency is modest, for the latter it is farmore significant. Diagrams have been constructed(Fig. 2) allowing us to evaluate the point at whichhydrates form under given operating conditions(temperature and pressure) as a function of the meanmolecular weight or density of the gas.

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methan

e

0.6 grav

ity ga

s

0.7

0.8

30 4040

60

80100

150

200

300

400

600

8001,000

1,500

3,000

4,000

6,000

50 60temperature (°F)

pres

sure

for

hyd

rate

for

mat

ion

(psi

a)

70 80 90

Fig. 2. Temperature-pressure diagram; hydrate formation curves(GPSA, Gas Processors Suppliers Association).

This fact is approximately valid since, as has alreadybeen said, hydrocarbons from normal butane to higherhydrocarbons play no part in the formation of hydrates.The same is true for inert gases such as nitrogen as,although these raise the molecular weight of the gascontaining them, they decrease their tendency to formhydrates. The most accurate way of evaluating hydrateformation is to calculate the vapour/solid equilibriumfor the gas mixture under examination. The mostcommonly used method is based on the equilibriumconstants Krs�y/x where y is the molar fraction of thehydrocarbon in the gas and x its molar fraction in thesolid phase. As for the determination of the dew point,the point at which hydrates form satisfies the followingcondition: Σ y/Krs�1. The values of Krs for thecomponents involved are determined experimentally asa function of pressure and temperature.

The care taken to avoid the formation of hydratesis motivated by the disastrous impact which theformation of solids in a gas transport pipeline and thepotential line plugging caused by this phenomenonmay have.

Eliminating the causes of this is extremely simple,and is based on the removal of water by dehydratingthe gas itself.

Inhibitors are used for short stretches of pipeline,for example the flow lines linking the well to thegathering and treatment centre where the dehydrationmentioned above is carried out.

Hydrate inhibitors are highly hygroscopiccompounds which can act both in the gaseous andliquid phase, such as methanol, CH3�OH,characterized by high volatility as well as by itshygroscopic qualities. Other inhibitors, by contrast,such as monoethylene glycol (C2H6O2) and diethyleneglycol (C4H10O3), characterized by very low volatility,act in the liquid phase.

The result obtained by using inhibitors is to lowerthe point at which hydrates form to a temperaturebelow the minimum temperature encountered duringtransport.

Suppose that we need to transport a gas through asubsea pipeline. The worst transport condition refers tothe temperature of the sea bed, assumed to be 6°C andthe maximum working pressure, assumed to be 90absolute bar. Let the gas be delivered at 90 absolutebar and at 20°C, and let it be saturated in waterupstream of the line; let the flow rate be 2·106 Sm3/d.For the sake of simplicity consider a gas consisting of100% methane.

According to the diagram in Fig. 2, the hydrateformation point of the gas is 54°F�12°C. To avoid theformation of hydrates in this case, it is thereforenecessary to obtain a value of 6°C. The diagram in Fig. 3shows the variation of the hydrate formation point when

EG (Ethylene Glycol) is used. In the case underexamination, the final acceptable concentration ofdiluted glycol is less than 50% by weight. To obtain theamount of glycol to be injected it is therefore necessaryto perform a simple material balance calculation. Theamount of water which condenses from the condition ofsaturation (at 20°C and 90 absolute bar) to the finalcondition (at 6°C and 90 absolute bar) is calculated: thisis equal to 16 kg/h. If the concentration of the glycolinjected is 85% by weight and the final concentration is50%, the amount injected must be no less than 25 kg/h.

When it reaches the end of the pipeline, the dilutedglycol is separated out and brought back to the desiredconcentration by regeneration. This is obtained byreconcentrating the glycol at atmospheric pressure byboiling off the condensed water. The glycol injectedacts in the liquid phase, and must therefore be mixedin an optimal way into the gas stream to be protected.In order to be effective, the glycol must wet the entiresurface of the pipeline; if the pipeline is very long, theamount of glycol accumulating inside may beextremely significant. This effect is accentuated by thebathymetric profile of the subsea pipeline.

In order to avoid significant accumulations ofliquid and the consequent reduction of the system’stransport capacity, it may be extremely useful todisplace the liquid by launching of spheres.Alternatively, an extremely large slug catcher must beinstalled at the end of the pipeline.

Finally, it should be noted that the reduction in thewater dew point obtained with this methodology issignificant. This value is slightly below the minimumtemperature of the pipeline (6°C) at arrival pressure,which is obviously lower than the initial pressure.

The gas thus obtained after the separation of thediluted EG does not meet transport and distributionspecifications (the water dew point required is below

686 ENCYCLOPAEDIA OF HYDROCARBONS

DEVELOPMENT PHASE OF HYDROCARBON FIELDS

100�40 0 40

50 % 25 % 0 %

temperature (°F)

pres

sure

(ps

ia)

80

1,000

experimental dataHammerschmidt

10,000

Fig. 3. Hydrate inhibition with ethylene glycol.

�10°C at 60 absolute bar), and it must thereforeundergo dehydration treatment.

It should be remembered that in the past, when gaswas transported from offshore platforms, glycol wassystematically injected.

A transport system thus involves the injection ofinhibitor after a preliminary separation at the wellhead,at a working pressure equal to the flowing pressure ofthe well. This solution makes it possible to avoid saltyformation water mixing with the glycol. The latter isinjected upstream of the pressure control installed onevery pipeline downstream of the wellhead separatormentioned above. In this way, the glycol inhibits thehydrates which might form by the cooling caused by thepressure drop in the valve described above (Joule-Thompson effect). The glycol injected is separatedonshore, regenerated, and sent back to the platformthrough a small transport pipeline coaxial with the maingas pipeline. This pipeline is installed together with thegas pipeline itself.

More recently, especially when the subsea pipelineis extremely long, an alternative solution, more simplefrom operating point of view, has been to install adehydration limit on the platform. In this way, when thenatural gas arrives onshore it can be sent straight intothe transport and distribution networks, although inmany cases the inhibitor injection system is still used.

As said earlier, methanol can be used instead of themore common EG to protect against hydrates when thereduction in the hydrate formation point required isimportant. Given an identical final concentration,methanol offers a lower hydrate formation point thanthat obtained with ethylene glycol.

The use of methanol for injection is not particularlycommon since a considerable amount of the liquidinjected passes in the vapour phase and is therefore lost;furthermore, glycol is less flammable and is neither

aggressive nor toxic, unlike methanol which has theadvantage of acting in the gaseous phase. It is thereforethe ideal solution for discontinuos use (for example toprotect a well during start-up or to melt a hydrate plugwhich has formed accidentally).

The above discussion explains why thediscontinuous injection of methanol is always coupledwith the more common continuous injection of EG.

Dehydration by coolingThe discussion in the preceding paragraph makes it

clear why one of the most simple ways of dehydratinga gas is to cool it, simultaneously injecting therequisite amount of inhibitor. The final temperature ofthe treatment is quite the same of the dew point whichwe wish to obtain.

In the case examined above, shall be the gas cooledto �10°C for example by expanding it from 90 to 60absolute bar.

Assuming the same composition (100% methane),from the enthalpy-pressure diagram for methane it canbe deduced that simple expansion causes a cooling ofabout 14.5°C. Therefore, to obtain the desired result, afeed effluent exchanger is required, allowing the gas tobe pre-cooled upstream of expansion from 20°C to 2°C. The gas at the outlet of the low temperatureseparator at �10°C will warm to about 9°C, coolingthe incoming gas (Fig. 4). The drop pressure assumedis therefore sufficient, coupled with a very modestexchange (or alternatively a more modest expansionwith a more significant exchange), to obtain thedesired result, in other words a water dew point equalto �10°C at 60 absolute bar.

Returning to the preceding example and theinjection of inhibitor, it can be seen that by injectingmethanol, a final concentration slightly below 50% byweight can be used, whereas for ethylene glycol an

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gas/gas exchanger

dry gas

to liquidgathering

system

EG

gas fromhigh pressureseparator LC

TIC

low

tem

pera

ture

sep

arat

or

PIC

Fig. 4. Conditioning gas with a lowtemperature separator.

amount such as to maintain a concentration of dilutedEG above 60% must be injected. In both cases,however, the amounts in question are extremely modest.

Now assume that a significant amount of EG isinjected (for example 1,000 kg/h with aconcentration of 97% in weight). In this case, theamount of water to be inhibited will be greater thanin the preceding case: about 23 kg/h. The dilutionresulting from the removal of water to the requireddew point is negligeable, and the glycol thereforemaintains a high concentration (95% in weight).This has an equally significant impact on thewater/gas equilibrium. From the EG/waterequilibrium diagram (Fig. 5), it can be seen that withthe concentration of glycol described above, therequired water dew point is obtained with a contacttemperature of about 15°C, much higher than the�10°C calculated previously.

Following the scheme described above, it istherefore sufficient to cool the gas to a slightly lower

temperature and inject 1,000 kg/h of concentrated EG.Since the cooling is very modest, the pressure droprequired to obtain it is equally modest.

To complete the analysis developed above, it shouldbe noted that, given an identical required water content,a higher pressure corresponds to a higher dew point; forexample, a pressure of 75 absolute bar instead of 60 inthe low temperature separator (LTS) corresponds to adew point of about �8°C, which increases the operatingmargin with respect to the required dehydration.

In many cases it is impossible to obtain a low costcooling of gas such as that described in the exampleabove. Under these conditions, the properties of glycoldescribed above can be exploited without the supportof cooling.

Dehydration by absorption with glycolThe glycol most frequently used for this type of

dehydration is triethylene glycol (TEG), but in somecases diethylene glycol (DEG) and monoethyleneglycol (EG) may also be used successfully; these havedifferent properties, as shown in Table 1.

Fig. 6 shows the scheme of the plant. A stream ofconcentrated TEG is fed into the top of an absorptioncolumn operating under pressure, into the bottom ofwhich the gas to be treated is injected. The contactwith glycol obtained with several equilibrium stages incounterflow with the gas leads to the dehydration ofthe latter as it exits the top of the column in question.The stream of TEG, diluted by the water which it hasabsorbed (rich TEG), is discharged from the bottom ofthe column. In a second column, operating atatmospheric pressure and provided with a bottom

688 ENCYCLOPAEDIA OF HYDROCARBONS

DEVELOPMENT PHASE OF HYDROCARBON FIELDS

Table 1. Physical properties of glycols

EG DEG TEG

Chemical formula C2H6O2 C4H10O3 C6H14O4

Molecular weight 62.1 106.1 150.2 Boiling point (°C)

at 1 atm 197.3 244.8 285.5 Viscosity (cP) at 25°C 16.5 28.2 37.3 Viscosity (cP) at 60°C 4.7 7 8.8 Decomposition temperature (°C) 165 165 206.7

200

0

20

40

60

80

100

120

140

160

180

20 32040 60 80 100 120 140 160 180 200 220 240 260 280 300

wat

er v

apor

dew

poi

nt (

°F)

solution temperature (°F)

0 %30 %

50 %60 %

70 %75 %

80 %85 % 90 %

92 %

94 %95 %96 %

97 %

98 %

99 %ethylene glycol by weight

Fig. 5. Gas dew points over aqueous ethylene glycolsolutions temperature.

reboiler, the rich TEG is reconcentrated by distillingthe absorbed water.

The regeneration column is divided into twosections: an upper rectifying and a lower distillationsection, separated by the glycol feed. In order toreduce energy consumption for regeneration, theconcentrated TEG from the bottom of the reboiler isexchanged with the feed. In order to produce anadequate reflux, and lower the losses of TEG due toregeneration, a condenser is installed above the refluxsection. This usually consists of a helicoidal coil whichcreates a heat exchange between the cold glycol to beregenerated, flowing through the inside of the coilitself, and the vapours rising from the reflux section.

The rich glycol exiting the absorber is saturated inwater and in equilibrium with the gas stream to betreated under the operating conditions of the column;it therefore contains dissolved gas. To liberate this gas,allowing it to be used as fuel, a flash drum is installedupstream of the regeneration column, operating at lowpressure (from 3 to 7 absolute bar). This is usuallyinstalled downstream of the top condenser describedabove, allowing the TEG to undergo modestpreheating, and facilitating the aforementionedseparation by reducing viscosity (TEG is the mostdense and viscous of the three types of glycolmentioned above; see again Table 1).

If the gas treated has a negligible higherhydrocarbon content, the separator (flash drum) onlyneeds to degas the glycol; in this case it can be avertical two-phase separator with modest hold-uptime. When the gas treated is associated with crude oil,

and therefore rich in higher hydrocarbons (light ends),the rich TEG may also contain liquid hydrocarbons.The latter cause a considerable increase in the foamingtendency typical of the glycol-natural gas system, andmust therefore be removed as efficiently as possible.For this purpose, a three-phase horizontal separator isused, with long liquid hold-up times (20-30 minutes).

Downstream of the flash drum, the rich TEG isfiltered with a cartridge filter and an active carbonfilter. For low flow rates of glycol, both filters operateon the whole stream; for large flow rates, the activecarbon filter treats only part of the total (from 20 to50%). Downstream of filtration and heat exchange, therich glycol enters the regeneration column. The bottomreboiler is usually of the kettle type, and the column isflanged directly onto the upper part of the reboiler itself.

Various systems can be used for heating. The TEGregeneration temperature corresponds to that of areboiler operated slightly above 200°C, since at 207°C the TEG itself undergoes significant thermaldegradation. It is thus evident that the finalregeneration temperature control must be very preciseand effective; this is also true for DEG and EG, whichhave a far lower degradation temperature (165°C).

Bearing in mind that decomposition, thoughmodest, may occur on contact with the hot wall of thereboiler, which must necessarily be at a highertemperature, it is important to limit its temperature andkeep it at a constant value. This is done by reducingthe density of heat flow in the boiler to a very lowlevel, obviously attempting to keep it uniform over itsentire surface (12,000 kcal/hm2). Although the overall

689VOLUME I / EXPLORATION, PRODUCTION AND TRANSPORT

TREATMENT PLANTS FOR GAS PRODUCTION

leanglycol

richglycol

free liquid

filter

dry gas

flash gas

flas

h dr

um

water vapour

stil

l

reboiler

surge drum

glycol contactor

wetgas

inletscrubber

Fig. 6. Simplified scheme of dehydration with glycol.

heat tranfer coefficient is fairly high, since it is mainlywater which must be boiled, the temperature drop atthe wall remains high.

Many TEG regeneration systems use fired tubesand fume tubes immersed in the bath to be heated.This solution does not always meet the requirementsdescribed above; even if the aforementioned heat flowdensity parameter is met, the latter remains a meanvalue which varies considerably over the exchangesurface, and which has extremely high values near theburners. Using fired tubes and fume tubes also makesthe reboiler far more bulky than an indirect exchangeobtained with a shell and tubes exchanger, or anelectrical heater with armoured heating elements.

If, as is frequently the case for gas and oil fields,the production of electrical energy for the overall fieldusers has a modest cost, the latter solution presentsconsiderable benefits in terms of simplicity, reliability,compactness and, most importantly, the reduction ofboth operating and investment costs. Since the cost ofthe exchange surface is modest, the density of heatflow can be reduced to 10 kW/hm2 (equal to 8,600Kcal/Hm2), at minimal cost and maintainingconsiderable compactness. This value, coupled withthe control of the regeneration temperature and thesystem’s intrinsic guarantee that the temperature of theheating element will be kept perfectly uniform over itsentire surface, creates an optimal heating system.

Where available, condensing steam at mediumpressure (20 absolute bar) has similar advantages;however, this condition is not particularly frequent.

Finally, a hot oil heating system may also be used;however, this does not give the same results as theelectric resistance described above since. Furthermore,given its characteristcs, this system cannot ensure auniform wall temperature. Using it provides resultscomparable to those of fired tubes; however it isobviously far safer and more reliable, though moreexpensive.

The regenerated TEG passes into the feed-effluentexchanger, where it gives up most of its sensible heat.Downstream of this exchange, it is pumped toabsorption pressure, and sent into the top of theabsorber (contractor) through a final cooler. Often thiscooling is carried out by exchange with the treated gasexiting the absorber.

Dehydration treatment by absorption with glycol isthus a very simple continuous process. In order toensure that the unit functions well, a surge vessel forthe circulation pumps is also needed; this is connectedto the kettle reboiler which feeds it by gravity. Theoptimal hold-up time for the vessel is 20 minutes.

With a regeneration system like that described, basedon pure thermal regeneration, the maximumconcentration obtainable for TEG is about 98.8% in

weight. This value can provide acceptable dehydrationonly when the absorption column operates under normalconditions, in other words with a operating pressureabove 60 bar, and an ambient temperature of 30-35°C.

When the gas must be compressed for deliveryto the transport and distribution system, it ispreferable to place the TEG unit downstream ofcompression.

The diagram in Fig. 7 shows that the equilibriumdew point corresponding to 98.5% by weight for TEGat a contact temperature of 50°C is incompatible withnormal transport conditions. Partly this is because it isnecessary to maintain a positive difference inconcentration between the dehydrated gas and thecorresponding condition of equilibrium (driving force)to obtain the desired dehydration. The driving forceallows mass transfer between the gas stream and theTEG itself to take place. The higher the driving force,the lower the number of stages of equilibrium neededto obtain the desired result, and consequently thesmaller the absorption column needed to carry out theaforementioned stages of equilibrium.

The amount of circulating glycol also has asignificant impact on the number of stages requiredfor dehydration.

690 ENCYCLOPAEDIA OF HYDROCARBONS

DEVELOPMENT PHASE OF HYDROCARBON FIELDS

�65

�60

�55

�50

�45

�40

�35

�30

�25

�20

�15�10

�5

0

0 5 10 15 20 25 30 35 40 45 50 55 60 65 70 75 80 85

5

10

15

20

2530

35

40

45

gas

dew

poi

nt (

°C)

solution temperature (°C)

99.9599.999.899.7

99.5

99

98

97

969590

99.97

Fig. 7. Water-TEG equilibrium diagram. The dashed linerepresents the concentration of lean TEG normally producedin a regenerator operating at atmospheric pressure and 204°C (Gas conditioning and processing).

The above discussion shows that the concentrationof the glycol is the decisive parameter for dehydration.When operating conditions require it, it is thereforenecessary to regenerate the glycol more deeply.

Below, the three most common deep regenerationmethods are described. Using these methodologies,very high concentrations up to 99.98% in weight andeven above can be obtained.

The most commonly used system involves installinga gas stripping column (dryer) below the reboiler. Theglycol, concentrated thermally in the reboiler, is fed intothe top of the aforementioned column, which containsrandom packing to a predetermined height (from 800 to2,000 mm). The stripping gas, preheated to 200°C, isinjected into the lower part of the column. The gas, incounterflow with the liquid (TEG), creates a masstransfer between the two streams, removing the waterand thus completely dehydrating the glycol.

The results obtained are excellent, and requiremodest investment and operating costs; the strippingcolumn is extremely small, and the type of internals isthe cheapest and easiest to install. Operating costs,linked to the amount of gas used, depend on theconditions at which the latter can be disposed of. If itis used as fuel, costs are very modest, and this solutionis definitely to be preferred. If, on the other hand, thegas exiting the top of the regenerator cannot bedisposed of directly, but must be recompressed andtreated like the main gas stream, the use of strippinggas must be carefully evaluated.

A second method for obtaining high concentrationsand avoiding the thermal degradation of the glycolduring regeneration is to use a vacuum system. The

regeneration scheme is identical to that describedabove; the only difference lies in the fact that thesystem is kept at very low pressure (0.1 absolute bar)by using a suitable vacuum unit (for example a systemof two-stage ejectors with steam as a driving fluid, or,when this is not available, a liquid ring pump).

The binary equilibrium of the TEG-water and DEG-water system varies as a function of the pressureat which regeneration is carried out. In this context, itshould be noted that at 0.13 absolute bar, very highmolar concentrations correspond to far lowertemperatures than those which cause degradation. Inthis way, as well as avoiding the resulting thermaldegradation, this system allows for the use of heatingfluids at temperatures which are more favourable froman economic point of view (for example steam at lowpressure, which can thus be produced by heat recovery).

When the gas treatment plant as a whole usessteam, this system makes for considerable savings(extremely low costs for the vacuum unit and a verycompact shell and tubes reboiler).

A third system, which unlike the previous two ispatented, is the so-called Drizo system, which involvesthe use of stripping gas obtained by vapourizing liquidcompounds with a suitable heating coil. Thesecompounds are essentially aromatic hydrocarbons(toluene, xylene, etc.), normally present in modestconcentrations in gases associated with crude oils.These are absorbed by the glycol, which theyconcentrate, and are recovered during the regenerationphase by condensing the top vapours.

The condenser and corresponding three-phaseseparator allow the aforementioned aromatic

691VOLUME I / EXPLORATION, PRODUCTION AND TRANSPORT

TREATMENT PLANTS FOR GAS PRODUCTION

STRIPPING GAS DRIZO��R�

water vapourand stripping gas

refluxcondenser

heatsource

rich glycol

lean glycol

stripping gas

reboiler

rich glycol

refluxcondenser

heatsource

solventpump

stripping solvent

reboiler

reboiler

rich glycol

lean glycol

water

rich glycolstripper stripper

rich glycol rich glycol

stand pipe stand pipe

Fig. 8. Simplified schemes of TEG regeneration.

hydrocarbons to be separated in the liquid phase,simultaneously separating out the water produced byregeneration. In this way, since both fluids can becondensed at atmospheric pressure and at relatively hightemperature, it is easy and cheap to recycle the aromaticsinto the TEG stripping column by vapourizing them(Fig. 8). With this system, concentrations of 99.99% byweight can be reached for both TEG and DEG. Thissolution, coupled with an adequate number of stages ofequilibrium in the absorber, allows gas to be dehydratedto up to 1 ppm by weight of residual water.

The material used for glycol plants, except inunusual cases, is carbon steel both for the absorber andthe regeneration vessel. The column internals, on theother hand, are in AISI 304 or 316L.

Dehydration by adsorption with molecular sievesWhen we wish to obtain an almost total removal of

water (0.1 ppm residual content), a solid bedadsorption process can be used. Unlike the precedingtreatment, this process is semi-continuous.

The molecules of water and of some polarcontaminants (CO2, H2S and mercaptans) are adsorbedby a silica gel which acts as a molecular sieve,allowing the gas to pass through the bed unaltered andretaining the polar molecules mentioned above, andwater in particular, in its active centres with a bond ofpurely physical nature. For water the adsorptioncapacity is extremely high: 20% by weight during thefirst adsorption cycle. Subsequently, following thethermal cycles characterizing the regeneration of thebed, this capacity declines due to the progressivedegradation of the adsorbent material. At the end of abed’s life (on average three years) this capacity falls toabout 13% by weight. Therefore, in calculating thevolume of adsorbent needed to carry out the requireddehydration, a slightly lower parameter is used (12%).

To obtain a continuous dehydration of the gas,several adsorbent beds are needed (usually three, ofwhich two working and one in regeneration).

The adsorption of water is carried out by makingthe gas flow from the top to the bottom of theadsorption column (down flow).

Regeneration is carried out by interrupting theadsorption cycle before the bed becomes completelysaturated in water; regeneration gas is used forstripping, suitably heated to about 280°C.

The regeneration gas is made to flow in the oppositedirection to the adsorption gas (up flow), in order toensure the complete removal of the water adsorbed.

The mass transfer from the gas to the solid duringadsorption is facilitated by high pressure, whereas theinverse phenomenon during regeneration is facilitated byhigh temperature and/or low pressure. Usually in naturalgas treatments the pressure is not lowered for

regeneration, which occurs simply as an effect oftemperature; this means that extremely high temperaturevalues are required. These values represent the mainlimitation of the dehydration system under examination.Heating the regeneration gas requires a heat source atextremely high temperature.

In small plants, where the regeneration duty ismodest, armoured resistances can be used for heating.This allows an optimal control over the temperature ofthe gas itself, and that of the heating surface.

For large plants with higher duty, the system usedto heat the regeneration gas is based on radiantheaters; this creates direct contact between the gas andthe surface of the heater coils subjected to radiation.

The most commonly used type of heater is thevertical pipe still with burners on the bottom (seeChapter 5.3). Regeneration takes place in several stages:the first is the rapid heating of the bed; the second andmost important takes place at constant temperature andinvolves the removal of the water adsorbed. Whenregeneration is nearly complete, the temperature beginsto rise again until it reaches its maximum value, veryclose to that of the regeneration gas. At this point thestripping of water is complete. The next stage is to coolthe bed, which must be brought back to optimaladsorption conditions. The gas passing through the bedduring this stage is simply dehydrated gas, which hasobviously not been heated. At the end of this stage thesieve is ready for a new adsorption cycle.

The adsorbent generally used is granular, withgrains of spherical or cylindrical shape (extruded); themost common sizes are 1/8'' or 1/16''. The acceptablesuperficial velocity of the gas, different for the twotypes of granules, depends on the working pressure.

Using this criterion, the bed’s cross-sectional area offlow is sized. Its height is calculated so as to create thevolume needed to adsorb the required quantity of water(this in turn depends on the length of the cycle which wewish to obtain). The density of the adsorbent material isabout 700 kg/m3. As far as regeneration is concerned,the sizing criterion is based on a simple energy balance.The heat supplied by the regeneration gas must providefor the various components contributing to the heatingof the bed, and the desorption heat of the water. Theheating must also take into account the increase intemperature of the vessel itself, and of lost heat.

Frequently, in order to reduce the latter twocomponents, but above all to reduce the thermal stresson the vessel, the adsorber is insulated inside thevessel allowing its metal wall to be kept at atemperature between that of the bed and the outsidetemperature. The ratio of regeneration gas to treatedgas flow rates usually ranges from 5 to 10%.

Fig. 9 shows a scheme of a typical dehydration unit.As can be seen, the regeneration gas is cooled as it

692 ENCYCLOPAEDIA OF HYDROCARBONS

DEVELOPMENT PHASE OF HYDROCARBON FIELDS

exits the bed, and the water which condenses as aresult of this cooling is separated. Downstream of theseparator, the regeneration gas is recycled at the twocolumns working in the adsorption phase. Thisadsorption-regeneration scheme is completelyautomated, with a suitably programmed processoroperating shut off valves which regulate flow and thusthe various phases of the cycle.

Note that the scheme described can be used for thedehydration not only of natural gas, but also of lightliquid products such as LPG and condensatesrecovered by cooling the gas itself.

Condensate removalThe removal of condensates from gas has various

aims, primarily to render it transportable. Duringtransport and distribution, the gas must meetspecifications on water and hydrocarbon dew points.This objective usually requires a negligible variation ofthe original composition of the gas, achieved bychanging the content of only some components(heavier hydrocarbons).

The condensate recovered can be stabilized in asimple way, identical to that developed for crude oil.The difference between this stabilization and that ofcrude oil lies in the product obtained, which is farmore similar to a light gasoline.

The vapours produced during stabilization arecompressed from the working pressure of the stabilizer

(7-10 absolute bar) to the feed gas pressure of thecondensate removal unit itself.

Often the gas to be treated is extremely rich notonly in heavy hydrocarbons (C5+), but also in higherhydrocarbons (ethane, propane and butanes). A typicalexample of this situation is represented by gasesassociated with crude oil, or those deriving fromcondensed gas. Under these conditions it is not alwayssufficient to meet the hydrocarbon dew pointspecifications. Since the composition of the gas is notmodified significantly, the second objective ofconditioning, concerning the control of the calorificvalue and consequently the Wobbe index, is not alwaysmet.

To obtain this result, the recovery of higherhydrocarbons must usually be increased, reducing thepropane and butanes content markedly. This obviouslyrequires deeper treatment.

If the gas to be treated is that associated with crudeoil, the production of gas is modest compared to thatof crude (low GOR, Gas Oil Ratio). The production oflight condensate by condensate removal thusrepresents a negligible proportion of the total oilproduction. This makes it possible to stabilize the lightcondensate by mixing it with the crude oil production,without modifying the vapour pressure required for theoil itself. In this way, most of the butanes and higherhydrocarbons can be recovered without the need tocommercialize a third product.

693VOLUME I / EXPLORATION, PRODUCTION AND TRANSPORT

TREATMENT PLANTS FOR GAS PRODUCTION

// // //water

regeneration gas

regenerationgas

drygas

wetfeed gas

closed

open

cooler

rege

nera

tion

tow

er

adso

rpti

on to

wer

regen. gasheater

FRC

550°F

500-550°F

inle

t sep

arat

or

knoc

kout

sepa

rato

r

Fig. 9. Two-bed dehydration scheme.

By contrast, when the treatment concerns gasproduction from a gas condensate field, the solutiondescribed above is difficult to put into practice. A fieldof this type is characterized by a far higher GOR thanassociated gas, and thus by a very modest flow rate ofliquid separated from the gas. As a consequence, thelight condensate (deriving from condensate removal),rich in propane and butanes, cannot always be mixedwith the stream described above. Under theseconditions, it is necessary to produce a third product,LPG, which is a mixture of propane and butanes withmarginal contents of ethane (2-4% mol) and C5+ (1%in volume). In addition to these limitations oncomposition, LPG is a finished product and musttherefore also meet all the relevant salesspecifications.

When production is modest, LPG is stored inpurpose-built cylindrical horizontal tanks, underpressure (11 absolute bar) and at ambient temperature;for safety reasons these are installed underground. Forlarger productions, pressurized spheres with a capacityof up to 2,000 m3 are used.

For extremely high productions and long storagetimes, as in the case of transport by sea, LPG is storedat atmospheric pressure by cooling it. In this case, thetwo components (propane and butanes) are oftenstored separately; they can be mixed together later inthe desired proportions. This allows the bubble pointof the final product to be controlled as a function ofthe climatic conditions of the market at which it isaimed. By lowering the propane concentration, forexample to 30% in weight, a mixture suitable for hotclimates is obtained; higher contents are acceptable fortemperate or cold climates. Where used as feedstockfor the petrochemical industry, propane and butanesare commercialized in the pure state.

Depending on the various objectives describedabove, different types of condensate removal can becarried out, from the simplest (with the gas having ahydrocarbon dew point of �7 to �10°C), up torecoveries which also include ethane and not merelyLPG.

The simplest and most obvious condensate removalprocess is based on cooling the gas, which can beachieved in numerous ways. One is that describedearlier to obtain a simple dehydration, based on theself-cooling of the gas by expansion. The scheme ofthe plant is identical (see again Fig. 4); the onlydifference lies in the properties of the gas to be treated.In this case, the feed to the condensate removal unit isnot a dry gas but a rich gas. In the low temperatureseparator, LTS, two liquid phases are thus separated atthe bottom: one of water and glycol, and one ofhydrocarbons. Although the two phases differsignificantly in density, and the dispersing

hydrocarbon phase has very low viscosity, the LTS isusually only two-phase. The separation of dilutedglycol and condensed hydrocarbons is thereforecarried out in a downstream separator dedicated to thisservice. To improve the conditions of the latterseparation, the emulsion may be preheated.

Once the two liquid streams have been separated,the condensate is stabilized and the glycol isregenerated. In this situation, the flow rate of glycol isvery modest, since it is used only to inhibit hydratesand not for dehydration. Furthermore, theconcentration of the glycol to be injected is kept at alower level (70-85% in weight) to reduce its viscosity.For this operation, monoethylene glycol is used almostexclusively.

Usually, in low temperature gas/liquid separationmore sophisticated and more efficient drop removalsystems than the simple wire mesh pad are used. Thelatter, particularly, is unsuited due to its potentialobstruction (and consequent breakage) due to theformation of crystals of hydrates and solid paraffinresulting from a malfunction in the injection ofinhibitor. To eliminate these drawbacks a cycloneseparator is often used.

It is worth noting that the LTS treatment schemedescribed above is used far more frequently forcondensate removal than for simple dehydration,where it is limited to a few specific conditions. Thelowering of pressure and consequent cooling have anopposite effect as far as water is concerned, whereasabove the cricondentherm the condensation ofhydrocarbons is facilitated by a reduction of pressurein addition to the drop in temperature due toexpansion. This means that, when removingcondensate from a gas, the optimal operating pressureis determined by the cricondentherm itself, or a valueclose to it. The value of this parameter for natural gaslies in the range of 30-50 absolute bar, correspondingto pressure levels which are not optimal fordehydration.

If the gas to be treated is produced at higherpressures, the gas expansion can be exploited; togetherwith the reduction in temperature this facilitateshydrocarbons to condense. Treatment is followed bythe compression required to transport the gas (pressureof 70-80 absolute bar).

In other cases, such as for gas associated withcrude oil, the gas is produced at lower pressures, andmust therefore be compressed before treatment. Underthese conditions, keeping the operating pressure of theunits as far as possible at the values specified above, adifferent cooling system is needed.

This condition can be obtained with an externalmechanical cooling cycle. In this case the temperaturerequired by the low temperature separator is obtained

694 ENCYCLOPAEDIA OF HYDROCARBONS

DEVELOPMENT PHASE OF HYDROCARBON FIELDS

with an exchanger (chiller) in which the cooling fluidevaporates at low temperature, removing heat from thegas to be treated. The other components of thetreatment unit do not differ significantly from thescheme described earlier.

The peculiarity of this solution is that the pressureof the treated gas remains unaltered through all theunits. The total pressure losses in the gas/gas exchangeand in the chiller can be kept as low as 1-1.5 bar.

The various refrigeration units used to conditiongas, and for the low temperature storage of the liquidproducts derived from it (LPG and NGL, Natural GasLiquid), will be described in greater detail in Section5.4.6.

With mechanical refrigeration, far lowertemperatures than those required for the simpleconditioning of gas can be reached. For example,using propane as a cooling fluid, the evaporationtemperature can be kept at about �40°C; as a resultthe gas can be cooled to –36°C.

In this temperature range a genuine dehydration ofthe gas must be carried out before cooling, since thehydrate inhibition system described earlier has a verylimited range of use.

To avoid the condensation of water, dehydrationmust obviously guarantee that the gas has a dew pointtemperature lower than that to which it will be broughtby cooling. The type of dehydration chosen is tailoredto the dew point required.

For example, if an associated gas must be cooled to�35°C, we use dehydration with TEG, stringentlyregenerated (99.95% in weight) with stripping or in avacuum. In some cases, an even more stringentdehydration with molecular sieves may be adopted. Inthis case the light condensate obtained in the lowtemperature separator (LTS) is known as NGL. Thisliquid stream is saturated in light hydrocarbons, andmust be adequately fractionated forcommercialization.

In these cases, the most common scheme involvestwo liquid products: LPG and stabilized gasoline. Theliquid product from the bottom of the LTS is fed into afirst fractionation column, or deethanizer, whose topproduct is a gas consisting basically of methane andethane; the bottom product consists of propane andhigher hydrocarbons (C3+).

After the recovery of frigories with an exchange,the top gas from the deethanizer is mixed with themain gas stream. Since the column in questionoperates at a pressure of 25-30 absolute bar,recompression is needed to obtain this result. Thebottom product is then fed into the next column, wherethe two liquid products are fractionated. The topproduct from this column is LPG, and the bottomproduct stabilized gasoline.

In some cases, rather than simply recovering LPG,a deeper recovery of higher hydrocarbons includingethane (feedstock for the petrochemical industry) iscarried out.

In this instance, a very deep refrigeration process isrequired: to about �100°C. This final gas temperaturecan be obtained with a multi-stage refrigeration cycle.The unit most frequently used under these conditionsexploits the self-cooling of the gas itself, supportedwhere necessary with an external supply of frigoriesproduced with a mechanical refrigeration cycle(propane cycle). To reach the low temperaturesmentioned above, the expansion of the gas itself isexploited, with the production of work through anexpander. For a simple evaluation of this process, referto the enthalpy-pressure diagram for methane.

An expansion is carried out with the production ofwork, obtaining a far greater degree of cooling fromequivalent expansion. In the expander, an adiabaticefficiency of 85% is commonly obtained; as a resultthe cooling obtained does not differ significantly fromthe corresponding constant entropy transformation.

Normally the expansion turbine is fitted to theshaft of a centrifugal compressor, which recompressesthe treated gas. Using this solution, the temperature ofthe gas can be lowered significantly with a modestreduction in pressure.

5.4.5 Other treatments

Sweetening

Absorption with alkanolaminesThis process has the aim of removing the acid

gases (essentially CO2, H2S and COS) present innatural gas. Mercaptans, when present in quantitiesexceeding the permitted limit (20 ppm of totalsulphur), must also be removed from the gas.

Whereas the removal of carbon dioxide andhydrogen sulphide is obtained with a single treatment,mercaptans are generally removed separately. Thereare numerous natural gas sweetening processes.Without doubt, the most common is based onabsorption with an alkaline solution (pH 11-13). Thistype of gas washing is thus based on a neutralizationreaction between the weak acid to be removed (H2S,CO2) and a suitable basic reagent. This reaction mustbe reversible, in order to allow the reagent solution tobe regenerated.

Table 2 shows the properties of the reagents mostcommonly used in the natural gas industry(alkanolamines).

Another reagent used instead of alkanolamines,and which develops a very similar neutralization

695VOLUME I / EXPLORATION, PRODUCTION AND TRANSPORT

TREATMENT PLANTS FOR GAS PRODUCTION

reaction to the latter in the water phase is potassiumcarbonate (K2CO3).

The basic absorption and regeneration reactions ofamines in the presence of H2S and CO2 can berepresented as follows:Reactions of the primary amine

velocity of reaction

2NRH2�H2S ↔ (RNH3)2S instant

(RNH3)2S�H2S ↔ 2RNH3HS instant

2NRH2�H2O�CO2 ↔ (RNH3)2CO3 moderate

(RNH3)2CO3�H2O�CO2 ↔ 2RNH3HCO3 moderate

2NRH2�CO2 ↔ RNHCOONH3R moderate

Reactions of the tertiary aminevelocity of reaction

2R3N�H2S ↔ (R3NH)2S instant

(R3NH)2S�H2S ↔ 2R3NHHS instant

2R3N�H2O�CO2 ↔ (R3NH)2CO3 slow

(R3NH)2CO3�H2O�CO2 ↔ 2R3NHHCO3 slow

Since this is a liquid phase reaction, the kinetics ofabsorption depend on the velocity of reaction, thepassage of the acid gas from the gas to the liquid phaseand the corresponding reaction equilibrium.

Frequently, diffusion at the film between the vapourand liquid phases represents the critical element in theoverall absorption process, including the equilibriumreactions in the liquid phase mentioned above. For thisreason, when carrying out absorption with a reaction,this aspect must be analysed very carefully in order tosize the absorber optimally. In practical terms, it can bestated that in order to reduce the impact of resistance todiffusion at the gas-liquid film the contact surfacebetween the two phases must be increased as much aspossible whilst simultaneously thinning the liquid film;in other words, internals with a high surface to volumeratio (random or structured packing) must be used.

Given an identical absorber volume, this choice leadsto a marked reduction in the liquid hold-up time insidethe absorber.

As far as the progress of the reaction is concerned,the situation is the opposite: it is obvious that a modestliquid hold-up time does not facilitate the completionof the reaction. There is a marked difference in thereaction kinetics of hydrogen sulphide with the basicreagent as compared to that developed by the latterwith carbon dioxide, particularly in the case ofsweetening with ternary and quaternary amines. Atypical case is that of MDEA (methyl diethanolamine)which ensures a highly selective absorption in favourof H2S with respect to CO2.

The above discussion explains how selectivity canbe significantly improved whilst minimizing the hold-up time by choosing suitable internals for theabsorption column and sizing these correctly, and byselecting a suitable reagent.

The hydrogen sulphide absorption reaction, whichrequires less time, can be carried out with a smallernumber of stages of equilibrium in counterflow. Byexploiting this peculiarity, H2S can be almostcompletely removed, as required by the transport anddistribution specifications for gas, without completelyremoving CO2; except in specific cases, this can remainin the gas even at relatively high values.

By contrast, when a deep removal of carbondioxide is required, whilst using the same reagentscharacterized by high selectivity (for exampleMDEA), the unit must be designed with a considerableincrease in both the number of stages of equilibriumand the liquid hold-up time. For example, in the caseof absorption with valve trays, a large number of traysare used. Additionally, to increase their efficiency, thehold-up time of the liquid on the tray is increased withvarious expedients, for example by increasing theheight of the tray liquid level by a higher weir andusing a single split flow configuration where possible.When the amine flow rates are high (high concentrationof CO2 to be removed) the latter option may involve an

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Table 2. Physical properties of amines

MEA DEA MDEA SNPA-DEA

Chemical formula HOC2H4NH2 (HOC2H4)2NH2 (HOC2H4)2NCH3 (HOC2H4)2NH2Molecular weight 61.08 105.14 119.16 - Boiling point (°C) at 1 atm 170.5 246 247.22 - Viscosity (cP) at 20°C 24.1 350 10.3 (at 30°C) - Reaction heat with H2S (kcal/kg) 305-371.5 277.3-332.7 283.4 Reaction heat with CO2 (kcal/kg) 343.8-388.2 321.6-360.5 362.1 Specific acid gas absorption capacity (mol/mol) 0.33-0.4 0.35-0.65 0.6-0.8 0.72-1.02

Concentration interval (% in weight) 15-22 25-35 35-50 25-35

excessive increase in the diameter of the absorber; inthis case a double-split flow tray is used.

Fig. 10 shows the typical absorption/regenerationscheme. The acid gas is washed with the regeneratedand suitably cooled amine solution in the absorptioncolumn. Absorption is facilitated by low temperatures,in part because the reaction taking place is exothermic.

The rich solution from the bottom of the absorberis first sent to a flash drum where the effect of the lowpressure (3-7 absolute bar) leads to a separation of apart of the gas and the solution is then fed into the topof the regeneration column after having exchangedheat with the regenerated hot solution. Here the acidgas is removed by stripping in a vapour stream.Basically, the inverse desorption reaction is facilitatedby low pressure (1.1 absolute bar) and the hightemperature obtained with the bottom reboiler (120-132°C). Like glycol, amines are subject to thermaldegradation which limits the regeneration temperature;this limit is even more marked due to the need tocontrol corrosion. To limit the regenerationtemperature, very dilute solutions are used; from thispoint of view the most stable amine, allowing forparticularly high concentrations, is MDEA (see againTable 2). The heating fluid most commonly used in thereboiler is saturated steam at low pressure.

Absorption with potassium carbonateThis unit does not differ significantly from the

amine solution unit already described. In the case ofabsorption with a hot potassium carbonate solution,the cooling of the regenerated solution at entry to theabsorber can be avoided or kept to a minimum(Fig. 11). This process was widely used in the pastbecause, as can easily be seen, it reduces theconsumption of heat energy for regeneration and the

use of large exchange surfaces. A second benefit, inspecific cases such as that analysed below, is linked tothe type of gas to be treated.

Like all other solutions in contact with a gaseousphase, the basic solution needed for sweetening has amarked tendency to form foam which is stronglyaccentuated by the presence of a third phase consistingof liquid hydrocarbons; these may be entrained in thegas or produced as a result of condensation.

Sometimes, during the removal of acid gas in highconcentrations, oversaturation conditions may developinside the absorber, linked in part to the reduction ofan inert gas, such as carbon dioxide, and the resulting

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sweetgas

acidgas

feedgas

flashgas

condenser

leancooler

rich/leanexchanger

reboiler

Fig. 10. Amine washing.

sweet gasacid gas

feedgas lean solution

richsolution

steam

Fig. 11. Absorption with hot potassium carbonate.

increase in the concentration of heavy hydrocarbons(C5+) present in the feed gas. Under these conditions,the cooling of the circulating solution to ambienttemperature, typical of the amine process, may lead tooversaturation and consequently to the unwantedpresence of liquid hydrocarbons.

The hot carbonate process avoids this condition bykeeping the whole absorber at a high temperature (80-90°C). The main disadvantage of this solution islinked to the need to replace the water which saturatesthe treated gas at high temperatures. To reduce thisphenomenon whilst simultaneously improving theunit’s energy balance, it is common practice to install aheat exchange upstream of the absorber, between thefeed gas and the treated gas, with the recovery ofcondensed water from the latter. This solution has asecond objective which may be decisive whensweetening treatment is followed by dehydration, forexample with glycol. In this case the simple removalof entrained drops does not protect enough thesolution circulating in the downstream unit fromcontamination. In addition to being more efficient, thewet removal carried out with water condensationsignificantly dilutes the concentration of the basicreagent in the water entrained in the gas.

The original hot carbonate process has beenmodified and replaced with the more efficientBenfield and Eickmeyer processes (Eickmeyer, 1962),shown schematically in Fig. 12. These two treatmentsdiffer from the hot carbonate process mainly in thatthey foresee cooling the stream fed into the top of theabsorber, in addition to the use of patented additives toreduce the corrosiveness of the solution to beregenerated. These additives mainly make it possibleto maintain a higher concentration of the reagent in theaqueous solution, thus increasing treatment capacityand reducing the amount of fluid circulating in theplant. The scheme used in the latter processesguarantees not only the partial removal of carbondioxide (final concentration of CO2 in the gas below1% mol), but also a removal of hydrogen sulphidewhich meets gas distribution specifications.

The use of membranesSemipermeable membranes can also be used for

the partial removal of acid gases, mainly carbondioxide (CO2), from natural gas.

The unit concerned is based on a simple principle:a mixture of several components, in this case in thegaseous phase, can be separated into two differentstreams by exploiting the selective permeability ofsome polar compounds present in the gas (such aswater and CO2), with respect to hydrocarbons. In thisway, a natural gas stream with a very high carbondioxide content and saturated in water, can be treated

with a membrane unit to obtain the partial removal ofCO2 (for example 80% of the original content) and asimultaneous dehydration of the gas. The pressurizedstream to be treated (42-48 absolute bar) is firstfiltered to protect the separation membrane from thepresence of solid particles or hydrocarbon drops.

To prevent the formation of condensation, whichmay also be caused by the oversaturation in higherhydrocarbons resulting from the removal of the inertgas (CO2), the gas is preheated to a temperature about20°C above its dew point.

Downstream of this pretreatment, the membraneseparates the gas into two streams. The first stream,consisting mainly of hydrocarbons, is the treated gas;the carbon dioxide and water removed areconcentrated in the second stream of permeated gas.The latter is at a pressure close to atmosphericpressure (2-3 absolute bar) and contains a significantquantity of hydrocarbons, mainly methane.

The higher hydrocarbons (propane, butanes, etc.), onthe other hand, remain in the main stream; in this way, byremoving the inert gas, a significant enrichment of thehigher hydrocarbon content in the treated gas isobtained, easing separation and recovery with a LTS.

This unit is extremely compact, and thus suited tooffshore platforms. The main limitations of its use arelinked to the significant loss of hydrocarbons (about20%) in the effluent (permeated) stream and the veryhigh operating costs due to the frequent need toreplace the membrane forming the core of the unit.

All the sweetening plants described must be madeof suitable materials to resist acid corrosion in both the

698 ENCYCLOPAEDIA OF HYDROCARBONS

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Fig. 12. Split-flow process with potassium carbonate.

sweet gas

acid gas

feedgas lean solution

richsolution

steam

pressurized absorber and in the regeneration unit(National Association of Corrosion Engineers – NACEcertified steels). Specifically, the top of theregenerator and the corresponding condenser andreflux accumulator require the use of AISI 316L orduplex stainless steel.

Notes on sulphur recovery and tail gas washing treatments

When the acid gas removed contains hydrogensulphide, the latter must undergo further treatment toturn it into sulphur. This process, which will only bedealt with summarily here, requires a series of reactionstages making up the Claus reaction: 2H2S�SO2 ↔2H2O�2S�heat. The reaction is carried out on acatalytic bed (bauxite) with stoichiometric quantitiesof reagents. The stoichiometric ratio is obtained bypartially oxidizing the hydrogen sulphide with air toform SO2 in a thermal reactor at high temperature(1,100°C or above).

The reaction temperature is maintained by the heatobtained from the partial oxidation of H2S to formSO2; the reaction temperature is also linked to theamount of inert gases present in the acid gas. Whengas with a high CO2 content is sweetened, the lattermay significantly dilute the hydrogen sulphide in theacid gas, causing a reduction in the temperature of thethermal reactor to below the optimal value mentionedabove. This occurs when the concentration of H2S islower than 50% by mol.

If the gas to be treated has an H2S/CO2 ratio lowerthan one, the best way to reduce the costs of the SulphurRecovery Unit (SRU) is selective sweetening, with thetotal removal of hydrogen sulphide coupled with apartial absorption of CO2. With a treated gasconcentration of 2-3% in mol of CO2, the amount of acidgas to be treated in the SRU can be greatly reduced.

Before the nitrogen and carbon dioxide formingthe tail gas can be released into the atmosphere, itscontent in residual sulphur compounds must also bereduced. For this purpose, a Tail Gas Clean-up Unit(TGCU) is needed. The latter consists of ahydrogenation reactor to reduce the sulphurcompounds to H2S, followed by cooling and a sectionof acid gas enrichment in H2S. This section washes thegas with MDEA, allowing most of the hydrogensulphide present in the tail gas to be absorbed. Thus,by regenerating the rich solution, an acid gas richer inH2S is obtained; this stream is recycled into the feed tothe sulphur recovery unit together with the main feed.With this SRU/TGCU coupling, the emission into theatmosphere of sulphur compounds can be limited to0.1% of the total sulphur produced.

It is important to note that the presence of CO2 inthe acid gas, and consequently in the SRU tail gas,

leads to a significant burdening of the whole washingsystem described, as well as of the unit itself.Recycling the stream containing H2S from the TGCUcauses further dilution of the SRU feed, since this toois strongly diluted. Considering that both the recoveryof sulphur and the TGCU used for this purpose areextremely important in the desulphurization of gas, theimportance of selectivity in the sweetening of naturalgas also becomes apparent.

Notes on physical absorptionIn many cases, particularly when the concentration

of H2S and CO2 in the gas to be treated is very high, itmay be preferable to use a physical solvent to obtaintheir removal.

Unlike the processes described above, the use of aphysical solvent does not lead to a neutralizationreaction with the consequent production of heat. Itsregeneration is therefore carried out without thecontribution of heat, by simple flash separation withseveral stages at decreasing pressures down toatmospheric pressure. Often, to improve theregeneration of the solvent, the final flash stage isfollowed by a cold stripping column using sweet gas.This solution can be used to remove acid gas eitherpartially or totally.

In recent years, the widespread use ofdesulphurization units for all petroleum products hasled to the sulphur market becoming saturated, makingsulphur difficult to dispose of. To avoid this problem,the use of partial removal units like that described,installed at the wellhead, is becoming increasinglywidespread for the treatment of gas with a high H2Scontent associated with crude oil production. In thisway, rather than producing sulphur, the solventregeneration gases are compressed and mixed withsome of the associated gas, and reinjected into thereservoir.

Notes on the removal of hydrogen sulphide with oxidation processes

When a gas has a very modest H2S content(for example 0.02-0.05% in mol), it can be purifiedwithout effecting the carbon dioxide content. Thisgives 100% selectivity, that is impossible to obtainwith any of the solvents described above.

The oxidation processes used to remove H2S alsohave another important benefit: they allow elementarysulphur to be produced directly during the regenerationof the rich solution, in addition to obtaining coldregeneration by simple oxidation with air.

This type of treatment originated a long time ago:it derives from the first processes developed for thedesulphurization of coke gas. The hydrogen sulphidewas absorbed and oxidized to form sulphur on a bed of

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limonite, iron sesquioxide, by blowing a small volumeof air into the feed gas to provide the stoichiometricquantity of oxygen required for the oxidation reactionof H2S to sulphur. Given its intrinsic properties, thisprocess is discontinuous, and is not suitable for thetreatment of large flow rates of gas at high pressure,typical of the natural gas industry. Recently, processesbased on this reaction have been developed which usethe iron oxide catalyst, kept in an aqueous solutionwith organic chelating agents (LO-CAT and Sulferoxpatents). This makes it possible to obtain a continuousabsorption/regeneration process such as that used foramine washing.

The H2S present in the gas is absorbed andoxidized in solution in the pressurized absorber. Theabsorbing solution is very dilute, and thereforerequires an extremely large amount of circulatingliquid to remove modest quantities of H2S; this is themain limitation of the process under examination.

The absorbing solution is then regenerated atatmospheric pressure and ambient temperature byblowing in air which frees the elementary sulphur inthe solid phase; this sulphur can be extracted from theregenerated solution with a rotary filter. If thechemism of the reaction is kept suitably under control,the cakes of sulphur obtained by filtration and dryingcan be easily handled and transported.

As mentioned at the beginning, when we need totreat a gas with a low H2S content, the filtration,handling, storage and transport of sulphur do notpresent a problem. Considering that even the sulphurproduced by the Claus reaction, though extremely pure,is not easy to sell due to the surplus, there no longer isany disadvantage in producing sulphur partiallycontaminated by the solvent, as in the case described.

The removal of mercaptans

These compounds are very frequently present ingases with a high hydrogen sulphide content, though infar more modest quantities.

Although they have acidic properties, more markedfor methyl mercaptan than for ethyl mercaptan, they arenot completely removed during the sweetening treatmentby amine washing.

Taking into consideration the concentration valuespermitted for the commercialization of gas, it may insome cases be necessary to proceed with furthertreatment.

In fact, the thermodynamic behaviour of the twocompounds (volatility similar to that of butanes formethyl mercaptan and that of pentanes for ethylmercaptan), means that a significant reduction ofmercaptans can be obtained with the cooling of the gasfor condensate removal.

It is easy to understand that a high recovery of thebutanes present in a rich gas leads simultaneously toan almost total removal of ethyl mercaptan and a morelimited, though still significant, removal of methylmercaptan; as stated above, this is removed to a greaterextent by amine washing. As a consequence, in anoverall gas treatment system which involvessweetening followed by deep condensate removal, bothmercaptans are removed by the combined effect of thetwo treatments.

It is worth noting that when sweetening involveswashing with a physical solvent at high pressure, thelatter treatment entails a far more stringent removalthan that which can be obtained with thecorresponding amine wash, and usually in itselfguarantees that the required specification is met. Theoxidation treatment described above, on the otherhand, has no influence whatsoever on the removal ofmercaptans.

Any undesirable traces of the latter for specifictreatments, such as the liquefaction of the gas, areabsorbed and then removed completely by themolecular sieves used for dehydration. With theexception of this particular instance, there is normallyno need to resort to specific treatment to removemercaptans, since this is achieved by the treatmentprocess as a whole.

The removal of mercury

Many plants for the high recovery of NGL and inparticular liquefaction units, in addition to thegeneralized use of stainless steels with a high nickelcontent, also have exchangers, such as plate fins, madeof aluminium and its alloys. In the presence of tinyquantities of liquid mercury, this material causes thetwo elements to amalgamate, resulting in thedestruction of the equipment itself. It is thereforenecessary to preventively reduce any mercury present inthe gas to be treated to extremely low concentrations.

By analysing the volatility of this element, it isobvious that its presence in natural gas may reach highlevels under reservoir conditions (high temperatures).At ambient temperature the harmful effect of mercuryis not significant because under these conditions itremains in the gaseous phase, and therefore does notreact with the metal surfaces with which it comes intocontact. At low temperatures, if present in significantconcentrations, it reaches the condition of saturationand, in the resulting liquid phase, produces the effectsmentioned above, especially dangerous forcomponents in aluminium.

Other materials commonly used in the gas industrymay also be damaged by amalgam with liquidmercury. For these reasons, a mercury removal unit

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DEVELOPMENT PHASE OF HYDROCARBON FIELDS

must be used in all processes which require the gas tobe cooled. This consists of a reactor on a solid supportwhich reduces mercury to values of a few ppb, and isshaped in a similar way to an active carbon filter. As itpasses through the filter, the mercury present in thegas reacts to form sulphides (HgS) allowing it to berecovered and handled safely. This removal is usuallycarried out after dehydration (with glycol), and beforeall subsequent treatments.

5.4.6 Process equipment and special units

Separators and slug catcherFor liquid/gas separation, the most commonly used

equipments are identical to those described in Chapter5.3, with the difference that in this case the mostfrequent separator configuration is vertical.

The double-barrel separator is one of the fewinstances where a horizontal configuration is used forthe treatment of natural gas. The two tubes connectingthe two superimposed vessels (down comers) aresufficiently large to dispose rapidly of the liquid slugentering with the gas, leaving the upper vessel of theseparator empty; this thus acts exclusively as aseparator for the drops entrained in the gas. The lowerpart acts as a collector for the liquid; the gas separatingfrom it rises into the upper vessel through the twodown comers. This separation system can be installedat the end of a short pipeline with condensation and theconsequent accumulation of liquid.

When the pipeline is large, and the hold upbecomes extremely significant, it is necessary to installa slug catcher. The most commonly used (finger type)consists of a set of parallel horizontal pipes linked toone another in a comb configuration. Each element ofthis comb acts like the double-barrel separatordescribed above. By increasing the number of parallelcomponents, and using lengths in the order of 100 mfor the lower part, extremely large volumes of liquidcan be captured and adequately separated, with a bulkybut relatively simple and cheap equipment.

In order to ease the gathering and drainage of theseparated liquid the two parts, upper and lower, slopeslightly with respect to the entry point of the gas; theupper part slopes upwards, and the lower partdownwards.

Heat exchangersThe heat exchangers used for gas treatment range

from the most common shell and tubes exchangers, tothe more sophisticated plate fin exchangers and thecryogenic exchangers used for the deep recovery ofNGL or the liquefaction of natural gas.

In recent years, very compact exchangers havebeen introduced for gas treatment; these represent adevelopment of plate heat exchangers (PCHE, PlateCompact Heat Exchanger), and are commonly used,especially for offshore treatment.

The choice of the most suitable type of exchangerdepends essentially on operational criteria (the powerrequired and the high temperatures to be reached) andlogistical criteria (available space combined with theneed to keep weight to a minimum). Another particularlyimportant factor is the exchanger’s maintenancerequirements, linked mainly to the properties of thefluids to be treated and their fouling tendency.

Generally speaking, when the exchange involvesreservoir fluids with a significant suspended solidcontent and a high fouling tendency, it is advisable touse traditional shell and tubes exchangers of thefloating head type which allow the tube bundle to beremoved from the outer shell. This operation allowsthe outer surface of the tube bundle to be mechanicallycleaned, and any individual tubes which have beendamaged to be easily replaced.

When the fluids to be exchanged are both clean,such as gas downstream of a separator, or better still asingle treatment unit, like that used for dehydration,far more compact exchangers can be used, such asplate fin exchangers. Plate fin exchangers are made ofaluminium, allowing them to operate even at cryogenictemperatures (�160°C); however, this material is notsuitable for fluids with a significant suspended solidcontent (low resistance to erosion as compared tosteel). This type of exchanger makes it possible toobtain very high exchange coefficients, and largeexchange surfaces per unit volume.

Under identical conditions, the PCHE allows bulkto be reduced further. It is important to note that thesemore compact and lighter exchangers are far moreexpensive than conventional shell and tubesexchangers. Their use is therefore limited exclusivelyto those cases (offshore platforms) where it is essentialto keep the space occupied and weight to a minimum.

CompressorsCompressors are extremely important in the

treatment of gas, since they are of decisive economicsignificance. The choice of the optimal type is basedon two fundamental parameters: the effectivevolumetric flow rate at suction, and the deliverypressure.

It should be noted that at low pressures rotaryscrew compressors are widely used; these allowextremely high compression ratios to be obtained witha single stage. They are also widely used ascompressors for refrigerant fluid (see below) and asflash gas compressors.

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Centrifuge compressors are the most commonlyused in the natural gas industry; their only limitation isthe minimum intake flow rate. With a multi-stagecompression system with intermediate cooling it isnevertheless possible to cover almost all applicationsin gas treatment.

When flow rates are low and delivery pressureshigh, reciprocating compressors are used; however,their use is more limited because they require moremaintenance than an equivalent rotating machine. Inaddition, the alternating motion of the piston, onwhich compression is based, causes undesiredvibrations and alternating horizontal forces which aretransferred to the machine’s supports, and thence tothe structures on which it rests. This characteristicrepresents a notable disadvantage, especially onoffshore platforms.

The motor most commonly used to drivecompressors is the electric motor. This solution isespecially suitable in the case of centrifuge and rotatingscrew compressors, even of large capacity. Its limitationis linked to the way the electric energy is generated, andthe power required by individual machines. Althoughelectric motors of 5 MW and above can be used, for highpower it is preferable to couple the compressor directlyto a gas turbine. In contrast to the preceding solution, thelatter allows the capacity of the compressor to be easilyregulated by varying the number of revolutions of theturbine; this leads to a significant reduction in the powerabsorbed when running slowly.

The use of electric motors to drive reciprocatingcompressors is limited by the need to insert a speedreducer between the driver (3,000-3,600 rpm,revolutions per minute) and the compressor (400-900rpm). This leads to considerable bulkiness and asignificant waste of energy. Internal combustionengines are therefore frequently used; these can becoupled directly to the compressor since the number ofrevolutions is identical. Where flow rates are low tomedium, this solution is frequently used (1-2 MW).

For very high flow rates and medium-lowpressures, axial compressors are used; these are morewidely used for the compression of the refrigerantfluids involved in the liquefaction of natural gas thanfor the compression of the gas itself.

ExpandersThere are two types of commonly used expanders

(gas expansion turbines, used in cooling systems).The first, described as a radial impulse expander, is

used for almost all gas treatment purposes. One of thelimitations to be taken into consideration where thistype of machine is used is the expansion ratio, whichshould not exceed three; for higher expansion ratios itis preferable to use several expansion stages since this

allows to obtain far higher performance levels. Thesemachines are usually small compared to the steamturbines from which they derive; however, there areapplications using expanders with a power (duty) of upto 5 MW.

The second type of expanders (axial turbines) isused for higher powers (duties) and extremely largeflow rates of gas; in the natural gas industry, thisapplication is nevertheless not common. For low power(for example up to 3 MW), the expander has anextremely high number of revolutions (15,000-18,000rpm). Therefore, the direct coupling with the centrifugalcompressor which uses the work of expansion requiresan extremely compact and efficient machine. On theother hand it is necessary to include an equally efficientlubrification system common to the two machines.

It should be noted that the use of expanders to coolnatural gas is becoming widespread, largely replacingtraditional mechanical refrigeration units. The reasonsfor this trend are simplicity and above allcompactness, as explained above. In addition, as caneasily be seen, this equipment allows the use of powersupplied from outside the cycle to be reduced, and insome cases avoided completely. This makes thissolution far cheaper than equivalent solutions based onthe Joule-Thompson effect or on mechanicalrefrigeration.

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DEVELOPMENT PHASE OF HYDROCARBON FIELDS

PB,PC

PATB

TB

TB=TC

SC

SD

TA

TA TD

hLB hLA hVB h'VD hVD

A

B C

D' D

Dh Qcd

critical pointbubble point

curvedew point

curve

pres

sure

(kP

a)

enthalpy (kcal/kg)

Fig. 13. Mechanical cooling: A, simplified scheme; B, enthalpy-pressure diagram (GPSA).

AB

C D

expansion valve

evaporator

compressor

condenser

A

B

Cooling unitsAbove, it was seen that many processes in the

natural gas industry depend on refrigeration. To this itshould be added that the storage and transport ofliquefied petroleum gases (NGL) and liquefied naturalgas (LNG) are based on the adoption of variousmechanical refrigeration systems. Bearing this inmind, we can analyse the plants most commonly used.

The principle on which mechanical refrigeration isbased is illustrated schematically in Fig. 13. Therefrigeration cycle is carried out by condensing apressurized fluid with a cold source at ambienttemperature, usually cooling water or air. The fluid is thenexpanded from condensation pressure to evaporationpressure through a valve. In the evaporator (chiller), thecooling fluid evaporates, removing heat from the fluid tobe cooled. Since evaporation occurs at constant pressure,for a fluid consisting of a pure component (such aspropane) the transformation is isothermal, and the heatabsorbed is equal to the latent heat of evaporation. As itexits the evaporator, the fluid is sucked up by thecompressor and brought back to the pressure at which itwill be condensed and then returned to the liquid phase.Using the enthalpy/pressure diagram for the cooling fluidunder consideration, it is possible to calculate the enthalpyvariations in the three main stages: condensation,evaporation, compression. The expansion in the valvewhich ends the cycle is obviously a constant enthalpytransformation. The data essential to analyse the cycle areobtained from enthalpy variations. The transformation D-A and the corresponding DH represent the heatexchanged in the condenser, divided into two parts: thefirst for the cooling of the compressed gas until thecondensation temperature TA is reached, the second forthe total and isothermal condensation of the fluid. Thistransformation takes place at the constant pressure PA.

The cooling supplied by the fluid corresponds totransformation B-C. The refrigeration fluid, atevaporation pressure PB, is in a mixed phase followingexpansion A-B and the corresponding cooling from thecondensation temperature to the evaporation temperature,which occurs as an effect of the partial evaporation of thefluid itself. The transformation C-D represents thecompression of the fluid, and the difference in enthalpybetween these two points is the thermal equivalent of thecompression work required by the cycle. Note that pointD differs from the corresponding point D� relating to theequivalent isentropic transformation. For the sake ofsimplicity, this schematized system does not considerpressure losses in the condenser and evaporator, and inthe pipelines connecting the various equipments in thisunit. The most significant of these losses, that in thecondenser, does not usually exceed 0.3 bar. For thisreason, the schematic representation does not differsignificantly from operating conditions.

The choice of the most suitable refrigerant dependson the properties of the fluid to be used, and on aseries of practical evaluations which range from inloco availability to the type of installation concerned,to the safety regulations adopted, and so forth.

However, the most important factor in selecting arefrigerant fluid remains the evaporation temperature,which must be a few degrees below the processtemperature. Also of fundamental importance is thetype of condensation to be used, with air or withcooling water. This is because the condenser isoperated at a temperature a few degrees above that ofthe cooling water and at least ten degrees above that ofthe surrounding air. Except in unusual cases, usingwater instead of air allows compression power to besaved by lowering the aforementioned temperature andthe corresponding pressure.

The illustration provided refers to a single stagerefrigeration cycle. Very often, the evaporationtemperature required is extremely low; in this case, weresort to a multi-stage cycle. The intermediate flashallowing the fluid suctioned from the second stage tobe cooled is known as an economizer, since its useentails a considerable saving of energy; this savingincreases the lower the temperature to be reached. It isimportant to note that the latter is limited, mainly forsafety reasons, by the minimum suction pressure of thefirst compression stage wich shall be slightly abovethe atmospheric one.

Multi-stage and cascade refrigeration systemsIn the natural gas industry, mainly in the storage and

transport of NGL at atmospheric pressure, and still morein the transport of LNG (Liquefied Natural Gas),extremely low refrigeration temperatures are required.The classic scheme used to obtain these temperatures iscascade refrigeration, shown in Fig. 14. This is based on athree-stage propane cycle, which allows a temperature of�40°C to be reached at 1.08 absolute bar (the interstagepressures are 2.5 and 5.8 absolute bar respectively,corresponding to temperatures of �20 and �7°C). Thiscycle is coupled with a further two-stage cycle, usingethane as a cooling fluid, and allowing an evaporationtemperature Te��87°C to be reached in the first stage ata pressure of 1.1 absolute bar. The interstage pressure is3.6 absolute bar (T��62°C), and that of the condenser13 absolute bar, corresponding to a condensationtemperature Tc��34.5°C (see again Fig. 14). Thistemperature is obtained by evaporating the propane fromthe first cycle. With these heat levels in the variousstages, it is possible to cool the fluid concerned, forexample NGL with a high ethane concentration to bestored at atmospheric pressure. As can be seen from thescheme, the three propane stages carry out the first threerefrigeration steps. The third temperature step supplies

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the frigories needed to make the ethane condense. If thisprocedure is extended to a system with three fluidsevaporating in cascade, the necessary temperature levelto liquefy natural gas can be obtained. This processrequires a temperature of about �162°C for puremethane. For gases containing higher hydrocarbons, suchas gas associated with crude oil, the bubble point of theliquid may be slightly higher, depending on the contentof the latter.

5.4.7 Liquefaction of natural gas

The above discussion helps us to understand thefundamental principles of the liquefaction of natural

gas, although the multi-stage system described earlieris not always used.

Before describing the liquefaction process, it isnecessary to define the conditions under whichliquefaction is carried out, and the gas compositionconstraints which must be met for commercialization.

Usually, the working pressure for liquefaction islinked to the system used to gather gas from the variousfields. A variation in operating pressure of 5 barsignificantly modify the temperature level of the mostimportant part of the cycle, corresponding to thepressurized liquefaction under pressure of the naturalgas. A lower pressure level leads to a lowering of thecooling curve, consequently penalizing the wholeliquefaction cycle. Furthermore, that part of the ethylene

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Fig. 14. Multi-stage propane-ethane refrigeration system.

cycle which is at the lowest temperature is increased. Asa result, the reduction in operating pressure increases thepower required by the ethylene cycle, and consequentlythe consumption of fuel gas needed to produce it; thisrepresents the most significant item of operating cost forthis unit. The following sub-cooling of liquefied gasrequires a multi-stage methane cycle. This phasecorresponds to the sub-cooling of the liquid needed toavoid excessive evaporation of the liquefied gas duringlater expansion to atmospheric pressure. In thesubsequent flash expansion, further cooling occurs dueto the self-refrigeration of the liquefied gas, proportionalto the liquid evaporating. If flash gas is used as fuel gas,a maximum evaporation slightly below 8% of the liquidis acceptable; thanks to this the liquid reaches the finaltemperature of about �160°C required for storage andtransport at atmospheric pressure.

The latter cooling can be considered the fourth andfinal refrigeration cycle. If the natural gas to be treatedhas a high nitrogen content, this can be reducedsignificantly with the aforementioned flash. N2, whichis far more volatile than methane, will concentrate inthe evaporated gas phase, thus reducing its residualcontent in the liquid. This options allows transport andregasification costs at the LNG terminal to be loweredby reducing an inert gas.

For all other inert gases and gas contaminants,treatment must take place upstream of liquefaction.Many of these treatments have already been described;this discussion will therefore be limited to listing themin the order in which they are carried out.

Upon the arrival of the feed gas pipeline, any liquidsare separated out, and the separated gas is filtered.Subsequently, it is decarbonixed with MDEA. If the gascontains H2S, this unit also removes this contaminant.With this treatment, we can simultaneously obtain thealmost total removal of H2S, and a residual CO2 contentof 50 ppm. The next treatment is the total dehydrationof the gas to a residual water content of 0.1 ppm withmolecular sieves.

The treatment with sieves also leads to thecomplete removal of hydrogen sulphide and othersulphur compounds, such as mercaptans, not fullyremoved by the sweetening treatment with MDEAdescribed above. Any residual carbon dioxide left afterthe latter treatment can also be absorbed and removedin the sieves. If mercury is also present, it must beremoved upstream of gas cooling.

At this point, the gas may enter the cooling andliquefaction cycle.

It is important to note that a transport project forliquefied gas requires a very important investment.The life of the project must therefore be considerablylonger than the 20-25 years normally foreseen for thedevelopment of medium-sized gas fields. The sources

feeding the plant must be multiple, and staggered intime. Under these conditions, it is probable that theproperties of these sources will vary over time, passingfrom relatively rich gas to relatively poor gas, and viceversa. If the range of variation of the calorific value,and consequently of the Wobbe index, exceeds thelimits specified for final distribution, we may resort tofurther pretreatment, consisting of a stage ofconditioning.

Following the first cooling cycle, for example to �35°C in the propane cycle, a separation stage isinserted with the corresponding fractionation of theliquids. The methane is recycled into the main streamthrough a demethanizing column; the C2+ can befractionated and sold as NGL or fractionated further intoLPG and light gasoline. This allows the calorific valueand the Wobbe index to be reduced to the required values.

During the 1970s, and to an even greater extent inlater years, the liquefaction plants with cascaderefrigeration cycles described above were largelyreplaced by the MRF (Mixed Refrigerant Fluid) system;however, this does not significantly improve the overallthermodynamic efficiency of the refrigeration cycles,and thus the energy consumption required to liquefy thegas. The main advantage of this system is its greateroperational simplicity and flexibility. The cascaderefrigeration is based on a series of refrigeration steps atdecreasing temperatures which do not adapt well to thecooling curve; as was seen earlier, this changes as afunction of operating conditions (pressure) and thecomposition of the gas (which may vary over time). TheMRF system, by contrast, uses a mixture of severalrefrigerants of variable composition instead of a purecomponent, adapting the evaporation curve to that ofthe fluid to be cooled. In this way, refrigeration and thecorresponding liquefaction of the gas can be subdividedinto only two cycles with two different mixtures.

The scheme in Fig. 15 shows the characteristics ofthis process (Precooled Telarc Process). The systeminvolves two large compressors (axial compressorswith extremely high flow rates and adiabaticefficiency), which ensure both precooling andliquefaction. The heat exchange is subdivided intothree exchanges. The first cools the gas with a plateexchanger (cold box) until it reaches the temperatureat which NGL and pure products separate; the latterare used to produce the cooling mixtures (N2, methane, ethane, propane, butanes and C5+). Twolarge cryogenic exchangers (spool wound) liquefy thegas and undercool it at �148°C and 48 absolute bar.Downstream of liquefaction and pressurizedundercooling, a two-stage expansion is carried out, to3.5 absolute bar and then to atmospheric pressure. Thegases produced downstream heat exchanger arecompressed and used as fuel gases. In part, they are

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also cooled and liquefied alongside the evaporationgas (boil off) from storage facilities.

The spool wound exchanger has a large exchangesurface (up to 50,000 m2) despite its dimensions, andallows several streams to be exchanged simultaneouslywith the refrigeration fluid, which evaporates as it isintroduced at various points and various temperaturelevels (the installation is vertical, as the scheme inFig. 15 shows). With this type of exchangers,exchanges with very modest temperature differentialsare obtained allowing the power required for thecooling cycles to be saved.

No mention was made of utilities in the discussionof other plants, and specifically the generation of thepower needed to drive the compressors. In this case itis worth remembering that engines for the latterrequire several hundred MW per line.

A liquefaction plant usually consists of three or morelines, giving a total capacity of 30-40·106 Sm3/d or more.The power required for it to function is equivalent to thatof a large thermoelectric power station (900 MW). Incomparison to old plants, where the power to drive thecompressors was produced with steam turbines(thermodynamic performance of the cycle of about30%), more modern units can use combined cycles withgas turbines and the production of steam generated bytheir hot fuel gasses, which in turn drive a steam turbine(overall efficiency above 50%).

Driving large axial compressors instead of thecentrifugal ones used for small capacities (adiabaticefficiency greater than 85%) with a combined cycle

reduces energy consumption for liquefaction, whichvaries from values in the order of 12-15% to values ofabout 7-8%. This energy saving technology, derivingfrom developments in the generation of electric power,thus has positive impact for the liquefaction and theoverall transport costs of LNG.

Bibliography

API (American Petroleum Institute) (1982) Guide for pressure-relieving and depressurizing systems, API RecommendedPractice 521.

Benedict M. et al. (1951) An empirical equation forthermodynamic properties of light hydrocarbons and theirmixtures. Constants for twelve hydrocarbons, «ChemicalEngineering Progress», 47, 419-422.

Campbell J.M. et al. (1985-1992) Gas conditioning andprocessing, Norman (OK), Campbell Petroleum, 4v.

Katz D.L., Lee R.L. (1990) Natural gas engineering.Production and storage, New York-London, McGraw-Hill.

Maddox R.N. (1977) Gas and liquid sweetening, Norman(OK), Campbell Petroleum.

Reid R.D. et al. (1977) The properties of gases and liquids,New York-London, McGraw-Hill.

ReferencesEickmeyer A.G. (1962) Catalytic removal of CO2, «Chemical

Engineering Progress», 58, 89-91.

Romano BiancoScientific Consultant

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DEVELOPMENT PHASE OF HYDROCARBON FIELDS

make up

LPG

�260°F

50 psia

LNG

NG

heavies

Fig. 15. Telarc process with precooling and separation of heavy components.


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