It seems that fuel oil does not have the same meaning in Europe and USA. In USA fuel oil can mean any fuel from diesel (No 1) to heavy residues. In Greece and Europe fuel oil (mazout) means No 5 or 6 per ASTM D-396, called heavy fuel oils in USA. For instance, diesel oil in Europe is not included in fuel oils.

Please could someone explain the difference between the following:

1 .Internal reflux

2 .Circulating reflux

3 .Pumparound

Answers

A:Ralph Ragsdale, Ragsdale Refining Courses, ralph.ragsdale@att.net

Internal reflux is the flow rate inside the column of the liquid from one tray to the tray below. The other two terms are interchangeable and represent the flow rate in the external piping of the liquid drawn the column, through heat exchangers, and back into the column either above or below the draw-off location/elevation.

During Crude distillation unit start up activities, water travels from crude storage tank to crude tower when furnace outlet temperature was 172C. It caused crude tower trays to dislodge. What if level of crude tower remains high then flash zone, does level of crude tower have significant effect on tray dislodge? Our system is furnace operated crude tower.

The sudden large water vaporization increase the vapor flowrate thru column internals creating an uplift force that dislodge the trays above the flash zone. Trays above flash zone are specified to hold 1 psi of uplift (some designers may prefer 2 psi, or 1/2 psi), this apply only to 3-5 trays located above the flash zone.

If bottom crude level is high (above the flash zone), and water gets into the column, then large water bubbles rising thru the liquid will dislodge the trays

What quantity of steam is required in distillation column and side strippers per barrel of crude/products.

Very sorry. I should have stated:

Column bottoms 0.2 lbs/gal bottoms

side strippers 0.1 lbs/gal stripper bottoms, not per barrel

Stripping steam have an important influence over the yields .

A general rule of thumb is around 10 -15 lb/barrel of product, but that ratio, which could be lower or higher than the previous range, should be optimized at the field and will depend of product nature. During field tuning, it should be considered the condenser capacity and hydraulic capacity of main tower or sidestripper.

Related to atmospheric residue stripping , the main objective is to recover light fractions as diesel or AGO. The bad performance or poor design of atmospheric residue stripping section could be a problem for downstream vacuum tower. There are many cases where atmospheric residue stripping section has poor design, then poor performance at any

steam/residue ratio .

For the case of side products the objectives could be as flash point control and to recover light material.

If you are executing a new design, 10 lb/barrel could be used and then optimized based on products spec, condenser duty and percent of stripped product.

Also it is very important to consider that stripping steam rate will impact the sourwater production.

Column bottoms 0.2 lbs/bbl bottoms

side strippers 0.1 lbs/bbl stripper bottoms

If there is a good reason to sharpen the cut between Atm Gas Oil and Atm. Bottoms, the bottoms stripping steam rate can be a part of the optimization calculations. But the reason must be valid considering the overall refinery marg

When water travels with crude, through furnaces into the crude tower, what process indications reflects on process parameters/tower profile?

Furnace feed control valves will get open for lots of water with crude and in extreme case the furnace will trip for low flow.

If you know you have water passing through with your crude in addition to the answers already provided you would need to make sure your tower top temperature is kept higher otherwise you will experience sever corrosion in you top trays and overhead system. Calcium Chloride common in most crude combines with heat and water to form chloride salts and hydrolyze to form HCl (dew point ~220F).

In addition to what Sam indicated you would also see an increase in the valve opening on the overhead water from the accumulator. With a lot of water carryover to the crude column you might also see an increase in delta P across the trays due to the increased volume of vapor flow that occurs when all the water flashes. The amount of firing in the furnace to maintain outlet temperature would also increase with a lot of water carry over due to the high latent heat of the water and the increased vaporization that would occur when the steam formed reduces hydrocarbon partial pressure.

f it is only a little bit of water you may see addition water going overhead of the crude tower or preflash tower,if it is a lot of water the preheat circuit could exhibit high pressures reducing flow rates, or the crude tower ovhd will over pressure , potentially lifting its safeties, or dislodging the trays

The presence of water vapor reduces the temperature at which a hydrocarbon vaporizes. Therefore, if the water in the crude is removed, and if the furnace outlet and column draw temperatures are maintained, directionally the overhead and side product End Points will decrease. The quantity of vacuum column feed will increase. To offset those changes when the water is kept out of the crude, increase the temperatures.

Can you please advise some literature sources or design guidelines for Naphtha Stabiliser design.

The purpose of the column is to reduce the vapor pressure such that the naphtha can be stored in an API tank in lieu of a bullet or sphere. Using the target maximum vapor pressure allowed in the region, specify the quantity of butane and lighter to remain in the product. Allowing some pentanes and heavier to go to the gas plant is O.K. as they will be recovered there. That means that it is difficult to justify an intense fractionation in this column. Thus, the reflux ratio in the column need not be high. With that goal, use the simulation to produce the heat and material balance data.

Some refineries do not have this column. The unstabilized naphtha is routed directly to the naphtha splitter (without tankage) which separates the Isomerization unit feed from the Catalytic Reformer feed. The specifications for that column depend on whether benzene is to be destroyed or recovered for sale.

In the crude distillation unit, we face problem with Gas Oil colour. Any one have any idea to solve this problem or any one have seen like this in any refinery!?

The Color deterioration is due to the following reasons:

1 .Crude processed containing more Nitrogen content. Mostly it will be reflecting up to Kerosene section. In worst case Naphtha section too will be affected.

2 .Flooding in the section beneath your Gas oil section. This may be reflected in the pressure drop measurement. Reason shall be, more partial pressure of Gas oil in the flash zone which is a reflection of improper crude blending containing higher gas oil content or stretching the throughput beyond the design.

3 .Very high sensitive Crudes which can crack in the Crude Heaters can contribute to this cause.

4 .Heavy cut of gas oil at considerably lower temperature of Draw off may absorb moisture and it may turn bright(not black).

In all cases, we need to scrutinize the position and try to solve it as soon as possible. The best safe solution would be change the successful crude blend composition and run to retrieve the position.

Brown/Black coloration of gas oil is a problem frequently observed in CDU and can be resolved easily with proper efforts. Such coloration may be accompanied with increase in metal contents too like V or Ni in the gas oil. Care may be taken for the reason that the contaminated stream with increased metals may have adverse effect on downstream

Hydrotreating or conversion catalysts .

The coloration may be due to poor fractionation with improper split of the cuts. The reasons for poor fractionation may be flooding, dumping at specific locations particularly wash zone region. The problem could crop up due to mal-operation, improper design of internals leading to residue entrainment in gas oil.

The tests like delta P survey in the zone, gamma scanning can throw light on the roots of the problem. I agree with other respondents stating the probable causes like bottom stripping steam rate, damage to wash zone, distributor issues for wash bed, wash bed coking etc. I know of published case where the coloration was caused due to an issue in the downcomer clearance after a revamp.

As mentioned it is a common problem due to residue entrainment with the Gas Oil, it can be caused ether by low liquid rate going to the wash section or high vapor coming from the feed, also flooding in the stripping section or damage in the wash zone internals (cause by operational upset) can deteriorate the color of Gas Oil

You will need to performed a series of tests to determine what is affecting the Gas oil color, some tests are:

Reduce Gas Oil draw

Reduce bottom stripping steam rate

Reduce Fired Heater outlet temperature

Based on your unit configuration you may try other tests that will help you to improve the Gas Oil color

Is this colour issue due to haze? Steam stripping of AGO saturates AGO and can then can form haze due to condensation of water after the cooler on the rundown. Its a pretty stable emulsion with an IFT usually lower than 10 dynes/cm however a high efficiency Liquid Liquid coalescer can break the emulsion and prevent the need for tank settling. It's possible to get < 15 ppmv free water after the coalescer which will be C&B provided the rundown temperature is not too high.

Gas oil colouration is normally due to entrainment of residue from the wash bed which may be malfunctioning or the stages provided in the bed are insufficient or the distributor (normally spray type) is maldistributing. If the column loads and internals details are available better analysis of the causes can be made. In fact we have carried out gamma ray scanning of the wash bed and the distributor to find out exact cause.

It is likely either entrainment or a leaking gas oil vs crude heat exchanger that is leaking crude oil into the gas oil stream. A number of things can cause entrainment - increasing rate,

change in feed stock to one with a higher gas oil yield, an increase in vacuum, coking of the wash bed or inadequate wash oil rate.

It may be either poor distillation internals allowing entrainment of jet spraying OR it may be a leaking heat exchanger. Check the heat exchangers first.

Very common problem due to entrainment. Did you increase the cut point? The tray efficiency? If the gas oil is fed to hydrotreater, you may see quicker catalyst deactivation.

I'm working on a study to design a new control schematics for Crude Distillation Column Pressure Control. Any ideas for CDU pressure control strategies?

Crude unit pressure control is to be done by using Fuel gas in the over head reflux drums. Generally second reflux drum will be used for this pressure control. Fuel gas is injected to second reflux drum and pressure reliving is to be done by taking out excess vapors to flare.

I am trying to build a model to optimize the operation of Crude Desalter and study its effect on Crude Column Overhead Corrosion.

The major salts present in Crude are NaCl, MgCl2 & CaCl2; but in our laboratory we measure only Total Salt Content of Crude (before and after Desalter); we do not measure individual salt.

My queries are:

1 )How the individual salt affect Desalter performance and Crude Overhead corrosion

2 )Is it required to measure the individual salt's content in Crude?

3 )Can I assume some typical break-up of individual salt (Note that the type of crude we process changes very often).

Typically the relative concentrations of chloride salts in crude oil mimic that of sea water. My rule of thumb is that 70% of the chlorides are in the form of NaCl, 20% MgCl2 and 10% CaCl2. I don't believe there is a significant difference in the desaltability of the different salts, although I have never seen data to prove or disprove this. MgCl2 is the primary contributor to overhead corrosion because it almost all hydrolyzes to form HCl. CaCl2 hydrolyzes but to a much lesser extent than Mg. NaCl does not hydrolyze to any appreciable extent. If you look at some old hydrolysis vs. temperature curves available in the literature, you will find that at typical crude unit conditions that about 20% of the salts in the desalted crude could be expected to hydrolyze in the absence of caustic. Some literature indicates that besides temperature, nap acid can increase the % hydrolysis.

While there can be differences in the salt composition for crudes, the effort to try and figure it out is probably not worth it. From a practical stand point you can't speciate the salts present in crude oil. You can do elemental analysis or ion chromatography, but in any case you have to make assumptions about what the actual salts present are.

One handy thing you can do is track the % hydrolysis of the salts leaving your desalter. Calculate the lb/d of chlorides in the overhead system from the chloride concentration and the total water. Divide this by the lb/d of chlorides in the desalted crude to get the fraction of the salts that hydrolyzed.

CDU overhead system

To provide a CDU with top reflux there are several configurations. The two main ones are as follows

Configuration 1: Double condensation

Configuration 2: Top pumparound

In configuration 1, total overhead vapors are condensed in two condensers in series. In the first condenser part of vapor equal to the required top reflux is condensed, water is separated and hydrocarbon liquid is returned to column as top reflux. The remained vapors (this includes the top product plus non condensable vapors) is routed to the second condenser.

In configuration 2, the top reflux is provided by a top pumparound and the overhead vapors ( (this includes ONLY the top product plus non condensable vapors) are condensed in a condenser and no liquid is returned to the CDU.

For a refinery with capacity above 100,000 bpd what configuration is recommended? Considering in both case energy is used in top/feed exchangers network. Can Heat Integration cause to use one of these configurations? We know in configuration 2 some more stages are needed as we have added one more pumparound!

Additional:

Thank you for the answers! Don't you think, the double condensation configuration results in lower flowrate in top section? When I can remove water in the first condenser, why not to return the reflux in lower temperature! furthermore I think thermodynamically, configuration 1 is better than configuration 2. As Ralph stated, the latter also needs at least two/three more trays for top pumparound! What I am not sure is energy saving! it is believed that configuration 2 results in a better heat integration.

The first configuration that you mention is used to reduce water condensation inside the column, which will lead to corrosion, also it reduces the probability of water condensation in the overhead, all this allow to use less expensive metallurgy in the mentioned areas. You have to make sure that water condensation happens at the outlet of the first condenser so chlorides can be removed.

For the second configuration, pumparound return temperature will be lower, which could lead to corrosion, specially that pumparound return will probably have free water (due to water solubility being reduce as temperature is drop in the exchangers)

Both designs are effective in reducing corrosion in the top section of the column, compared with a design that uses a single overhead accumulator. In any of the 3 cases, resist the desire to yield light naphtha overhead and heavy naphtha as the first side draw product. In other words, in any case, the overhead vapor temperature needs to be well above the temperature at which water is a liquid.

An advantage of the top pumparound (PAR) only design is that the crude exchanger does not need to be elevated for “free draining”. This savings can directionally offset the cost of additional trays and column height associated with the top PAR only design.

One thing to compare: For the general power failure case, and using the “hot firebox, tube boilout criterion”, there can be a significant difference in the size of the relief system. It has to do with the “Cox Chart” effect on LMTD of the air and water cooling portion of the heat removal configuration when calculated at relieving pressure.

To what extent can we blend fuel oil into gas oil without affecting the viscosity characteristics and maintaining the flash point specifications for gas oil or to keep them

within the allowed limits?

Additional info:

First of all we don't have neither FCC, Hydrocracker nor VDU...we only run a conventional CDU

the objective here is to maximize the yield of gas oil...(we call it solar in our national markets) by extra stripping out from fuel oil or residue...the question is; Is there any equations or experimental methods to calculate or estimate the resulting viscosity and flash

point of either the gas oil or fuel oil?

Thanks a lot.

Gas oil is quality is governed by the 90% or 95% (T90 - T 95) recovery temperature (whatever is the applicable case for the refinery) as well as the cold properties viz., Pour point and cloud point. One blends the Fuel oil to the extent that both their recovery temperature and the cold properties are not violated.

From your query, it looks that you want to blend Fuel Oil (i.e. long residue or atm bottom in your case) into gas oil - Solar oil which I guess is diesel. The answer is pl do not do as metals and other CCR components in fuel oil will get added into diesel which is no desirable for high efficiency diesel automobiles. Further, the heavy boling distillates in fuel oil will generate substantial smoke while driving and end up in engine deposits too. The best way to lift the front end of long residue into diesel is to operate the Crude tower at the possible lowest pressure and higher furnace outlet temperature within possible lowest overflash.

To clear every one's doubt about the question, please spell out clearly whether it is GO to fuel oil or fuel oil to GO. This will help to give more appropriate answer for the question.

If you are referring to gas oil as conventional feed to the FCCU, many refiners increase the cut point of vacuum unit bottoms until the maximum coke burning capacity of the FCCU is reached.

If the gas oil is hydrocracker feed, increasing the End Point is not recommended due to the affect on run length.

If you mean “gasoil”, aka “diesel fuel”, the amount of resid shifted to the gasoil can be determined in the lab. Distill the front end of the resid under vacuum as done in a crude assay, collecting small cut samples. Make a series of blends of gasoil and these cuts until the blend no longer meets the gasoil specifications when tested.

If you mean dropping diesel fuel into FCCU feed, some refineries have had to do this when the diesel fuel market is depressed. If you do it incrementally by cut point, lab test blending will tell you when the diesel fuel will still meet specifications. Otherwise, dump full range diesel into FCCU feed until you reach a capacity limitation in the FCCU, or until your LP model tells you that it is a mistake to run that operation.

Can you please advice what type of corrosion inhibitor, biocide, antifoulant and polyelectrolyte polymer can be used in Desalter effluent?

We have had success with merus rings to solve various problems including biofouling and corrosion. The ring is fitted to feed pipe. For info: www.merusaustralia.com.au

Quality of Desalter effluent Brine (pH, Oil content) is very important as brine joins the inlet streams of Waste water treatment. Brine pH could be lower while processing opportunity crudes. In such cases to protect the brine piping from the Desalter to WWTF corrosion inhibitors (simple pH boosters) shall be used. However, before dosing any chemical in Desalter effluent brine impact on the downstream sections must be thoroughly validated.

A chemical injection into the desalter effluent is not very common. Sometimes a cationic charged polymer is dosed into the desalter effluent to separate oil from the water at the API.

Beside the injection of emulsion breakers it is more common to inject metals removal additives and corrosion inhibitors into the wash water in front of the desalter. In that case the effluent water will contain most of these metals. Antifoulants like asphaltene dispersants are often dosed into the crude oil after electrical desalting.

If the desalter is performing properly you should not need all of the mentioned chemicals. What is the fate of Desalter effluent?

We have a very strange problem, it's that the desalter outlet crude has greater salt content than that of the inlet... the lab examinations proved that more than once...this always happens when the injection water is cut off-while switching from a tank to another. What

could explain this?

Additional info/response:

1 .We cut off water while switching between tanks because of the existing water accompanied with the crude from the new tank; I mean the first 30 minutes after switching to a tank, the crude has too much water to inject more.

2 .How could the NaOH type could affect this situation?

At the outset, I congratulate the person for asking a very good question which received such good response. My observations are as follows:

1 .Some of the refiners temporarily stop make up wash water when they switch over the crude tank. It is to prevent water shot in the desalter and subsequently to CDU due to large amount of water coming with the crude.

2 .Under such situations it becomes important to increase the dose of demulsifier appreciably very high. Monitoring and controlling of the interface layer and water drain valve becomes very important parameters to follow and act

3 .Agar Corporation have invented very good design for such situation and control of desalter operation. It can be read at their site namely http://www.agarcorp.com/Animation/Desalter.html

4 .Number of experts have expressed their opinion on use of caustic pre / post desalter. It is useful to overcome the issue of naphthenic acid. Sodium naphthenate formed in desalter make emulsion very tough to break.

Well, for a stable desalter operation, its better to have continous injection of water. The reason is 1. if teh inlet water content is low, injection water will actually improve the dehydration to certain extent and thus will have better desalting as well. when injection water is cut off with low inlet water, dehyration/desalting may not be as efficient as it should be and also at this point of time the actual water carry over should be checked as it could be higher as well. The reason for having higher outlet salt than inlet could be because of doing the test method wrongly. for instance, if the inlet BS&W is 0.3% taken at 15.6C aand outlet sample is taken at 120C and cooeld it to 100C and not 15.6C will surely result in high water content & high salt content as water will be more soluble at higher temp.

The solution is to drain the water from crude tank and then only to send it for desalting, without cutting the wash water. This will ensure that the desalter functions properly.

When you feed crude without draining the water, you are loading your desalter with crude containing brine which is already saturated with salts and therefore desaled crude will have same salt contents as feed crude.

Adding caustic will not solve the problem,use of caustic in desalter is resortedfor high acid crudes and is done after a thorough analysis of the effects on down stream processing.

Also please check the quality of water you are using for desalting, clean and low TDS preferred if using Stripped sour water ensure proper pH as per recommendation of Desalter supplier.

If problem pesists consult desalter vendors

Just cool the crude sample before you put in the sample bottle. What is happening when the crude is hot and not enough cooling it vaporize and you are concentrating the salt in the bottle. you think that you collect a litter for example to fill in the bottle but actually, you collect 2 litters , one vaporize and one remain in liquid phase , at this moment the salt in the bottle is not for one litter it is doubled and finally you find the salt content in the outlet is

higher than inlet ..

I experienced this issue myself and solved it by the proper cooling the sample before you put in the bottle.

Cutting off water as part of preparing for the Crude tank switch can be avoided by proper mixing, settling and draining of Crude tank before the tank is lined up for unit feeding. Opportunity / heavy crudes call for longer times of mixing as the water at times remain as pockets in the startified layers of crude blends. To avoid this stratification one can store neat crudes in separate tanks (an expensive option) and carry out inline blending if the Refinery

needs to process a blend of opportunity crudes so as to maximise profits .

NaOH injection will only help in improving the pH of Crudes when the pH is lower than the recommended range of 6.5 to 7.5. It can help in neutralising chlorides such as Mg Cl2, Na Cl2 that can lead to high Chloride corrosion. One needs to be very careful with NaOH injection as it can lead to a variety of extremely difficult to handle, downstream issues viz., poor desalting and poor dehydration when the crude pH goes beyond 7.5, preheat fouling due to poor dehydration, Catalyst poisoning in FCCs, Premature coking Cokers etc.

First I would like to ask about NaOH injection point up stream or down stream the desalter

injection of NaOH up stream the desalter cause formation of Na Napthenes especially if you processed high napthenes crude which act as emulsion stablizer and this may be the cause of the problem

your problem also may be due to over mixing across the mixing valve so try to establish optimum pressure drop across mixing valve (as if pressure drop is higher than optimum it will cause water carry over and higher salt outlet)

No situation will warrant to cut off desalting water while changing the feed tank. Once you are sure that the tank farm operations are taken care to prepare the feed tank well, that is makiing the tank free of water before feeding the unit, then you need not cut the desalting water. Even you feel water carryover in crude after feed tank is changing, you can cut off water after you realise water in crude. The trick is in preparing the feed tank. Once you cut

off water it is obvious that salt carry over is natural .

NaOH may not affect the desalting process adversly, it helps only to maintain the pH of brine (desalter drain). But it will contribute to increase the metal content in VGO and this will be poison for the down stream units lile FCCu/HCU.

Another possible source of the problem could be the sludge carry-over, with related emulsion (which is far tighter than the rag layer at the interface). You could evaluate online cleaning the desalter. ITW has developed and patented a technology for doing the job in as low as 24 hours on a oil-to-oil basis.

If you inject NaOH that could contribute to the result you see. Are you using D3230 for analysing the salt?

Why do you cut the wash water, as I understand, completely, why not reduce it 40-50%.

If you could analyse the constituents by AAS/ICP I would assume this could help guiding you to what is contributing to your higher values

I guess my first question is why are you cutting off the desalter wash water when you switch tanks... this is not a common practice and probably contributing to your problem... one possible cause for what you are seeing is that when you cut the washwater off the desalting process is stopped and you may be losing some of the desalter emulsion into the outlet which could then increase the amount of salt coming in

What is the best way of judging the efficiency of a desalter?

Desalters are the Key equipment in CDUs and does the functions of Desalting as well as Dehdyration. There are a few parameters that indicate the efficienty of Desalters, viz., Desalting efficiency, Dehydration efficiency, Brine quality, desalting water consumption, chemical consumption. While desalting helps minimising corrosion, dehydration helps in preventing fouling of down stream preheat exchangers (improve preheat recovery and hence minimise the Fuel consumption) and Brine quality helps in ensuring effluent quality in Waste water treatment facility at a low expense. Desalting is said to be efficient when 90% of Desalting is done in each stage (target of <1.0 PTB salt and <500 PPM of Oil in Brine). NO water carry over into Desalted Crude is the other parameter that judges the desalter performance. All the above need to be achieved with lowest Desalting water and Chemical consumption. Remember water is not cheap. High oil carry over with Desalter Brine increases slops production and hence increase recycling costs.

The foremost objective of the desalter is to remove inorganic salt such as chlorides of sodium and magnesium from the crude oil. Also it acts as a surge drum to give sufficient residence time to help settling of suspended particles. It should remove all water including micro size particles of water using the high electric field.

Efficiency is calculated as follows: (Quantity of salt removed/Total salt in crude before desalter)*100.

The expected removal percentage is 95 to 98%. Removal of water can be identified by closely observing the variation in sour water make from atmospheric column overhead product drum or reflux drum.

Efficiency of desalter can be improved by adding correct quantity of water (3 to 5% of feed), maintaining differential pressure across the mix valve (approx 1.0kg/cm2), desalter

temperature to improve solubility and desalter operating pressure to give proper residence time for settling of water. Desalter temperature also plays a role for settling of water by reducing the viscosity of crude oil. As the settling velocity is inversely proportional to viscosity, it helps to settle water.

Very good indicators for the desalting performance are the desalting efficiency (%) and dehydration efficiency (%). A single stage desalter system should give a desalting efficiency >90%. With a double stage desalter system a desalting efficiency >98% is expected. The

inorganic salts of NaCl, MgCl2 and CaCl2 are typically expressed as ppm NaCl .

The desalting efficiency is calculated with this formula: Desalting efficiency = ((ppm NaCl in crude oil before desalter – ppm NaCl in crude oil after desalter) / ppm NaCl NaCl in crude oil before desalter) * 100

To calculate the dehydration efficiency the amount of wash water must be included: Dehydration efficiency =((%water in crude oil before desalter + %wash water to desalter - %water in crude oil after desalter) / (%water in crude oil before desalter + %wash water to desalter)) * 100

In order to best judge the 'efficiency' of a desalting unit (whatever efficiency means) one compares the quality of the product with the feed, and estimates that improvement compared to the maximum theoretically possible. In the practical world, one measures the un-removed contaminants in the de-salter product stream and the oil lost in the de-salter water product. Salt content in pounds (kilograms) per thousand barrels is a commonly used metric. Also important is the BS&W to determine how much dispersed water is not being removed. Then a Karl-Fischer water analysis (ppm) to determine the dissolved water

remaining in the processed crude is helpful .

The process objective of 'de-salting' is to :

a) remove water soluble chlorides and sulfates and any other 'salts' so they do not cause corrosion as they decompose and form mineral acids in the distillation equipment

b) remove any particulates such as sand, iron oxides, clays, and any other 'sediment' that would accumulate on the distillation trays or packing and cause both plugging and corrosion

under the deposits .

The goal is a perfectly 'clean' feed to the distillation equipment. 'Efficiency' can be considered as a calculation of just how far short of perfection the unit and operations achieved.

One might also be concerned about the quantity of chemicals used, electricity consumed, pumping horsepower, water use versus theoretical minimum for the measured performance, hydrocarbon loss in the water effluent, total treating cost per barrel, etc. The BIG MONEY savings are in reducing residual salts and sediments as much as reasonably possible to reduce corrosion and fouling of the equipment.

KW per Barrel

%salt removal per stage

oil in brine

Required transformer size varies with crude source, so care is required in comparing with other desalters.

The simplest way is to measure the residual salt, i.e. 'Salt = ~ 1 pound per thousand barrels is good target for single stage desalting of conventional crude oils, while 'Salt' = ~0.5 ptb is

good target for recommended two stage desalter system .

Other aspects to consider are:

(a )Residual water in crude oil [should be down in the few tenth weight percent via BS&W testing]

(b )Residual dirt and rust and any carbonaceous materials in crude oil [Most of these should have collected in the bottom of the desalter vessels and been blown down to suitable external decanting and transfer receiver via well engineered 'mud drain system'

(c )whether emulsions have been mostly 'broken' such that there is only a small 'rag' layer inside the desalter vessels [with reasonable expenditures on demulsification chemicals]

(d )Ensuring that desalting is at optimal crude oil outlet temperature year around [regardless of whether crude oil in tankage is cold and encon exchanger network is clean or fouled].

What makes outside shapes of distillation columns differ from one another? i.e. shape of pre-flash differs from CDU, CDU differs from VDU?

I concur with the responses by Keith and Ralph.

The key drivers for cross-sectional area are the ACFS,vapors or Am3/sec, vapors and the feasible / comfortable capacity factors of the internals [trays or structured packing or grid packing] and the design margins. Vessel cost and structural considerations favor only a limited number of diameters in any given tower. Sometimes the bottom swage down below

the flash zone is internal in crude atmospheric towers .

Best practice designs for crude atmospheric towers often favor high liquid wetting flux pump-arounds using structured packings [as Lesson Learned from complex distillation tower theory / computer modeling]. In these instances, the hot gpm per ft2 or hot Am3/hr-m2 can become an important factor in determining cross-sectional area in conjunction with ACFS, vapors or Am3/sec, vapours.

Distillation column cross sectional area necessary to meet design thruput is determined by the volume of vapor and liquid that must be processed on that 'tray.'

A preflash column is sized to allow sufficient disengaging space for the light ends so 'flashed'.

A flashed crude column will typically have a larger mid-section since much of the 'work' of distillation is done in that section of the column. The top end may be either smaller, or

larger, depending on operating pressure and the light-ends yield of the design crude. The bottom of the so-called Atmospheric column will be sized to accommodate the work necessary to separate the gas-oil from the reduced crude.

A vacuum distillation column is similarly designed but has a much larger cross section because of the low density of the vapor due to the very low absolute pressure. Often the very bottom section is much smaller because both the vapor content and liquid rate are so much lower that the reduced crude feed and residence time at the typical high temperature must be minimized to avoid thermal degradation that will form deposits in the column and

plug it.

Sometime the physical strength requirements are such the very bottom section is too small to support the column and, and a 'skirt' must be added around the bottom section to provide necessary strength

Usually due to vapor rates being much lower in some sections, such as below the flash zone of the atmospheric column and the vacuum column. Save the cost if a smaller diameter is needed in that section. Another example: In a reboiled absorber, the vapor rate in the absorber section, above the feed, is much lower and can have a smaller diameter.

if the column pressure increases, will the separation increase, and vice versa?

Increasing pressure makes the separation worse directionally. That's because k values of adjacent components converge as pressure increases.

We want to by-pass our de-salters in order to check the consequences with and without desalters on CDU. Moreover we have stopped de-emusifier dosing prior to desalters. What impacts are anticipated in your opinion and what parameters to be monitored in case when

there is no desalter in crude preheat trains?

1 .A desalter de-salts; that is eliminates part of the CaCl2 and MgCl2 occluded in the crude. If you do not do this, well, both of this salts (none the NaCl) decomposes with the water and heat (120 - 370 °C) to HCL, hydrochloric acid, which is "mortal" to the walls of the CDU mainly at the head - coolest - zone (water condensation). In the other hand, a desalter helps to decrease the water and solids/dirt content in the crude. If you do not do so, you will get troubles on the heated zones of the furnace and the CDU (fouling, over boiling, erosion and corrosion)

2 .A demulsifier de-emulsify; that is to break the emulsion crude-water. If you do not do this you will increase dramatically the water content in the crude; the result: instability in the CDU operation (over boiling)

Consequences of bypassing desalter are already well known. What prompts you to do so? Demulsifiers do great job in water separation

Answer to question lies in “why desalting is done”. Crude coming to the refinery contains water, some free and some as emulsion. The water also has dissolved salts in it and possibly suspended solids, depending on source of crude. In desalter the water along with salts are

knocked out. Not using desalter is not a good practice and cannot be recommended. Literature deals in detail about it and should be consulted.

In case non-desalted crude is used as feed in CDU, serious problems can be encountered depending on quantity and quality water coming with the crude. Anticipated major problems are:

1 .Process upsets

2 .Fouling of exchanger train, erosion of heater tubes and transfer line by dissolved / suspended salts

3 .Upset and lack of consistency in the O/H corrosion control

All these in long run would damage equipment and result in loss of production and high cost of maintenance.

Answer to what parameters to be monitored in case when there is no desalter in crude preheat trains are:

1 .By-pass desalter only if it is required for maintenance

2 .Ensure max. drainage of water from the feed crude tank bottom before it is used as feed,

3 .Monitor out let and inlet temperatures of feed at critical locations in the feed exchanger train to see if desired degree of temperature rise is achieved or not. This would indicate extent of fouling, if any, and also identify affected exchangers.

4 .Operation would have to be vigilant to avoid upsets in distillation column in case of entry of large quantity of water .

5 .As there would be large variations (fluctuations) in HCl going in the O/H, the O/H corrosion control system cannot be manually operated and advanced automatic system will have to be installed.

6 .Any increase in erosion / corrosion of heater tube and transfer lines should be monitored.

Perhaps the crude you process is extremely light, clean, and does not have significant chloride salts.

However, the industry trend is toward significantly increasing the performance of their de-salter operations .

Dirt (fine particles from the producing formation), water-usually a complex brine, rust particles, and other miscellaneous deleterious matter normally MUST be removed from the crude oil before being heated to above about 200C. This is the approximate temperature at which the chloride salts dissolved in the inevitable brine in the crude begin to decompose into insoluble deposits and chlorine -which immediately turns into hydrochloric acid--which has a ferocious appetite for steel and its alloys.

The entrained dirt and rust particles will drop out of suspension, depositing in all the quiet regions on the trays, packing, horizontal surfaces,etc. plugging them up You may easily calculate the volume and mass of these potential deposits from the BS&W analysis of the crude. If those are not removed by the desalter, they WILL remain inside the heat exchangers and distillation tower.

IF the crude you process is similar to MOST, and you cease desalting, I would expect to find much more rapid decline in heat transfer coefficients on the pre-heat exchangers, requiring more severe firing of the furnace--resulting in more coke formation and higher pressure drop there, greatly increased corrosion rates on all the metals in the top section of the tower and all the associated equipment (HCL is not picky about what it eats up), and within a few weeks/months, the fractionation performance of the tower itself will have noticeable decline, requiring higher reflux ratios to maintain product distillation specifications. All of these effects significantly lower the operating efficiency of the unit.

Excellent answers, in my opinion. I would add the obvious: If the salt in the crude is 1 lb/1000 bbls or less, the desalter is needed only to remove BS&W, which in itself can cause fouling, but is not as critical as the corrosive effect of higher salt levels.

1 .If you bypass desalters, salt,sediments and water content can't be separated from crude.You can't add wash water to crude if desalter is bypassed and salts will carryover to preheat train and foul the exchangers. Caustic and magneesium salts will decompose above 200 C forms Hydrochloric Acid in Column overhead system which in turn increases corrion in overhead system.

2 .If you stop de-emulsifier, water in emulsion phase will not be separated even with Electric field. De-emulsifier reduces the surface tension between crude and water and helps to separate water from crude.

3 .If Desalter bypassed, post desalter preheat train fouling will increase and pressdrop across each exchanger increases which can be observed at last exchangers (>200 C). CDU colum overhead system pH will drastically drop to 3 to 4 because of HCl in aqueous form and rate of corrosion will be increased. Overhead chlorides will be increased which is a direct indicates of salts carryover in crude to preheat train and heater.

you should verify the following parameters:

Chlorides content in overhead water draw

Iron content in overhead water draw (to try to measure an increase in corrosion)

Sodium in the residium (If sodium if above 20 ppm it will speed up the coke formation in the fired heater of the vacuum column)

monitored the fouling on the heat exchangers downstream (this will change within a long period of time)

If you have a coker or visbreaker unit, also monitored the sodium in the feed to these units.

Bypassing Desalters is a BAD idea within the context of CDU+VDU operations. This will cause accelerated fouling of the hotter crude preheat exchangers plus salt fouling and accelerated coking within the Crude Heater plus accelerated deposition + fouling + corrosion of the top trays [assuming that CDU is not best practice with "double overhead drum system" or "high

recirculation warm TTPA and single overhead drum system ."

Desalters are intended to water wash the crude oil to remove dirt and rust and particulates and as well as lower the salt levels [both in emulsified brine form and micro-particulates].

The only plausible reason to bypass a desalter---except temporarily to fix or upgrade its internals or to retrofit a second stage desalter---is that the API Gravity of the crude mix is less than 19 degrees. Even then, best practice calls for recycling overhead naphtha or TTPA naphtha to bring the API Gravity seen by the desalter to 22 degrees or higher.

this is not a very wise move in my opinion, besides the potential increase in fouling of not only your preheat train but also your furnace due to salts and solids you can count on increase in corrosion concerns in the top of your cdu tower... suggest that you review your operations and ask yourself:

1 )what am I trying to achieve, by shutting down the desalter, are the financial gains from shutting down the desalter at the CDU offset from all of the other downstream cost

impacts?

2 )is the desalter working to its fullest potential, has it been optimized?

Besides the immediate impact on preheat, furnace and tower over corrosion you will see much larger losses in energy and due to the impact anticipated on downstream processes (if they exist), for example a slight increase in water can increase the energy loss by >$500 K USD/per year; increases in salt can result in FCCU catalyst impacts that can be >$1 M USD; there are similar impacts on thermal conversion units also.

What is the relationship between smoke point and ASTM D-86 of kerosene? Does an increase in IBP increase the smoke point or is it the increase in final boiling that increases the

smoke point?

Just to fix ideas in case the OP wonders why more aromatics means lower smoke point...

Aromatics (contains unsaturated benzene rings) have higher carbon-to-hydrogen ratio. And this is that deficiency of hydrogen that can cause unburned HC (smoke) during combustion in a Jet engine.

More aromatics means lower smoke point (low is bad).

Highest the FBP, poorest the Smoke Point but that depends on the molecularity of the FBP. Remember that the ASTM D 1655 – 09 establish a minimum Smoke point of 25 mm, or a minimum of 18 mm and a Naphthalene content of 3.0 vol, % maximum. Naphthalene is located at the FBP tail.

Smoke point is directly related to the Aromatic, if you done more aromatic saturation, then the smoke point will increase.

A good simple correlation to estimate smoke point from TBP and gravity data is probably that published by (Albahri, T. A., Riazi, M. R., Alqattan, A. A., "Analysis of Quality of the Petroleum Fuel," Energy & Fuels, 2003, Vol. 17, pp. 689-693.). The relationship therein is that increasing API gravity and Mean Average Boiling Point increase the smoke point. So increasing your IBP or FBP may reduce API gravity more than the benefit of adjusting the

MeABP, which might reduce the smoke point .

Mr. Ragsdale is correct that increasing the FBP is more likely to increase the aromatics content, which is bad for smoke point.

I'm new to the refining field. We have eletrostatic desalter with three transformers (each of 22KV). and I have questions about this:

1 .Why is a single heavier transformer is not used instead of three?

2 .We have 2 PSVs on desalter. What is the reason for two PSVs?

3 .Why the water injection before desalter is maintained 4.5% of crude charging?

4 .What is the function of grills (grids) inside the desalter?

1 )Three phase supply - three transformers , one on each phase balances the load. Also if one breaks, you still have some treatment.

2 )Not sure - what are the set points and design pressure? Are there two separate PSVs with different set points? Are you sure its a PSV andnot something else ?

3 )Ensures the dilution is sufficient to reduce the salt content to spec

4 )The grids are connected to the high voltage outlets of the transformers, the oil passes between them, the water droplets become polarised, coalesce and travel towards the grids and are removed. Try google books for a longer explanation.

1 .Desalter manufacturer will have to answer the first question. My guess is that it's so that if a transformer fails, you don't lose all the electric field.

2 .I don't believe that there is a fundamental reason why a desalter would need 2 PSVs. If it's a large vessel, you might have a very large fire load so that the required load exceeds the capacity of a single PSV or there could be a large blocked in liquid relief load if the vessel is not designed for the blocked in head of the charge pump.

3 .Wash water is added to enhance the desalting by improving contacting of salts with water, diluting chlorides in water, and enhancing water coalescing (more water means water droplets are closer together, which makes it easier for them to coalesce & settle). 4-5% is a common amount of water for lighter crudes. If you are processing a heavy crude you might need 7-10% wash water to help with the desalting.

4 .The grids create the electrical field used to coalesce the water droplets. The size of the transformer and the spacing between the grids set how strong of electric field can be generated for coalescing the water.

One of the observations pertaining to 2 stage desalters is that 2nd stage efficiency remains poor as compare to 1st stage (50 to 60 % in 2nd stage as compare to 91 % in 1st stage) despite trying all necessary parameter adjustments such as Delta P, Fresh water flow, interface level, emulsion checking etc.

Is there a means of remedying this discrepancy?

The question I would have is what is the actual salt number going to the 2nd stage? if it is low then the relative removal percentage may not be 90%, but something lower. If the salt going to the 2nd stage is >10 % then it would be reasonable to expect >90%. There are ways to calculate if it has reached the ideal level by determining the mixing efficiency (function of dehydration and salt removal) as well as determining the optimal salt content (function of the dehydration and inlet salt conc.)