7 replies to this topic

[afdmello]

Propylene is stored as a saturated liquid at ambient conditions of around 77F in winter and then pumped at 261psig via a level control valve to a tank.

The level control valve at ground level . the small tank 30 ft above metering pump suction.The tank is insulated. The pressure in the tank is 130psig and the TI on the tank indicates a temp of 75F.

The propylene if saturated should have a temp of 67 F corresponding to 130psig isnt it. Why is the pressure in the tank dropping to 130psig.

The tank feeds the metering pump ,and has globe valves in the suction line can we replace them with gate valves? the flow rate of C3H6 is around 660 lb /hr.

I am a chemical plant operator seeking information .

thanking you

AFD

[Art Montemayor]

AFD:

I answered this query two days ago, just before the Forum was changed over. I don't know if my response resided sufficiently for you to read through my reply, but I'll try to re-write it as best as I can remember.

Firstly, the basic data seems to have changed a bit from the original entry. No matter, my comments and suggestions still apply.

I made the following comments:

1) You don't state what the pressure in your Propylene storage tank is, so I have to assume it is the corresponding saturation temperature at 77 oF, 177 psig (approx.). I refer to the GPSA Databook Pressure-Enthalpy Chart.

2) I have to presume, like before, that you are lifting the liquid Propylene to a hydrostatic height of 30 feet to obtain a favorale NPSH for your metering pump (Pump #2); I also assumed that you are controlling the level in the Overhead tank with a level detetector that controls the strorage tank transfer pump (Pump #1).

3) If you merely want to get the high NPSH, why don't you eliminate any throttling on pump #1 by simply having an equalizer-overflow line between the Storage and the Overhead tank? I am attaching an Excel flow diagram that clearly depicts what I am describing of your installation. When you throttle the saturated liquid Propylene just for control purposes, you are inducing a flashing situation and creating a 2-phase problem downstream. You want to avoid flashing and 2-phase flow unless that is what you want to achieve in the first place. Also, when you throttle this application, you must relieve the pump's high pressure by recyling the discharge or you build up high forces on the pump's seal and induce overheating and vaporization with subsequent loss of prime. Why go through all of this?

4) When you equalize both tanks, the thermodynamic conditions are the same in both vessels - except for the hydrostatic height. We are only talking about 300 kg/hr (approx. 2.5 gpm) of liquid Propylene as I recall, so why not run the transfer pump continuously or, if don't want to do this, control it with a start-stop switch actuated by your level control? The equalizer line remains.

5) Your thermometers may be uncalibrated or they may be inserted in the vapor space of the vessels - which makes the temperature reading lag and is an unaccurate way to detect temperature in a 2-phase vessel. I would always put the thermometer in the liquid phase.

You said you had 3 globe valves in the suction to the metering pump and I can't figure or justify that type and that many valves. This is a simple application of a Full-Bore Ball Valve (in your case, about 1-1/2" to 2" in size). Your flows are so small, you have to rely on mechanical stability and that is why I opt for that size valve. Don't forget to specify that the ball be of the Trunnion design. This ensure seat and seal integrity. Also do not fail to specify that the ball be drilled with a 1/8" hole in the upstream side in order to vent when the valve is in the closed condition. If you don't do that with ball valves in high vapor pressure liquids, the trapped liquid inside the ball starts to expand and exert pressure as it heats when the valve is in the blocked position. This exerts a tremendous pressure on the ball's seats and supports and makes the ball relieve through seat failure (thank God!).

I would not use globe or gates on this service.

I hope this addressed your concerns and questions and helps out.

[afdmello]

Thank you Art. I was honestly expecting ,as is normal,a detailed reply from you.

The Propylene is produced at another place far away,stored at ambient conditions(77F) and is transferred by the first pump to us (users)and the pressure at our side is 261psig.The storage tank pump is continously on and I dont know what is the set up there.The LCV is controlling only the tank level and opens/shuts only on level no interlocks with any pump.

The tank 2 as stated by u is elevated for NPSH in our unit.

The Globe valves 2 nos for isolation of suction strainer and 1 for the pump .

Art the tank PT is indicating 130.5psig (depends on weather as summer goes to 174F)means that the liquid must and should be at its corresponding saturation temp of 68F isnt it??assuming the PT is correct and the TI may be faulty .If the TI is correct and in the vapor space should it also not show the the saturation temp. of 68F.The tank is insulated and having a capacity of around 50ft3.I checked the TI is in liquid phase but may be faulty.

Art,in this scenario,if the upstream pressure is 261psig the downstream presssure is 130psig(it varies a little from day to night)because of the pressure drop across the vavle OR flash cooling of propylene?.Actually what causes this pressure to drop to 130psig ?what can be made not to drop it so much ( we cant float the tank with supply as tank having PSV for vapor think)or what can make the pressure drop further.

Why wouldnt u use gate valves?I cant find the Excel flow diagram.

Just academic interest making me curious on these phenomenon I observe and not got any answers from my people.I could not catch ur reply in the other forum.

Fortunately for me I had saved all the precious replies you had given for my previous queries.Thanks for your time.My Profound thanks to Phil and you.

AFD

[Art Montemayor]

AFD:

First, let me apologize. I was probably the first one to use the new Forum format and I tried to take advantage of what I thought was a way to finally include attachments to our responses, like Excel sketches or calculations. I took the time and prepared an Excel flow diagram depicting what I think is your process flow, complete with data and calcs. But when I sent the response, I was advised in an error box that the file format could not be handled; so the attachment idea failed. Anyway, let me address your rightful concerns:

So, you are receiving saturated liquid Propylene at your battery limits at a pressure of 261 psig & 75 oF. You are feeding this liquid to an elevated feed tank according to level demand in the tank, by using a throttling control valve. If the elevated feed tank is at 130 psig and the liquid Propylene is saturated, the corresponding temperature that it is at is: 50 oF. Note that this is 18 oF lower than your estimate of 68 oF. As I said, I am using the values from the GPSA Databook, which I think are accurate enough for what we are estimating.

I wouldn't worry about a discrepancy up to 5 oF between our figures. But 18 oF is quite a difference. Anyway, I don't believe that you are worried about the absolute temperature value as much as you are about the situation regarding your ability to successfully maintain a positive NPSH and your operation producing on schedule as it normally should without any hazards or process upsets. And I agree with this. Allow me to analyze your questions using what I now know of your process:

1. Yes, the temperature in the vapor space is the same temperature as in the liquid. However, the vapor space has a thermal conductivity coefficient that is very poor (most gases and vapors have this characteristic) compared to that of the liquid phase. Many times, there is a lag in the temperature read-out due to the film coefficient and the heat leak to the surrounding atmosphere that takes place, and this cause erroneous readings. Since your vessel internal temperature is around 50 oF and your external atmosphere is at around 75 oF, you are experiencing a heat leak that affects your TI and makes it read on the high side. That's why the liquid phase is favored as the source of the temperature reading. If the heat leak is excessive, internal thermocouples are employed to obtain an accurate read-out.

2. Your pressure of 130 psig in the feed tank as compared to the available 260 psig in the incoming feed is due to the throttling of the liquid control valve. Something is controlling the vapor space pressure of the feed tank and since I don't have your P&ID and you haven't identified how the pressure is established in the feed tank, I can't tell what it is that is controlling your 130 psig vapor space pressure. I can only guess, and I think that you probably are bleeding down the vapor space in the feed tank to a lower pressure vessel or other part of your process, through a back-pressure control valve or a regulator. This is the nomal way a process engineer would do it if we are unable to equalize the pressure with the outside liquid Propylene source tank (as I have previously explained). Now, I think, you can appreciate why I highly recommend an equalization between both tanks and a simple liquid transfer. Any time you throttle a saturated liquid, you will create a "flashed" 2-phase flow, composed of cooled (and saturated) liquid and vapor at a lower pressure. The vapor, if left to accumulate in the feed drum will soon equalize with the feed 260 psig pressure and the flow of liquid propylene across the control valve will cease. In effect, what has occurred is that you have dead-ended the flow. In order to continue flowing liquid Propylene into the feed tank, you must vent or evacuate the flash vapors being formed by the throttling control valve. Again, this is why I favor equalization between both tanks; any excess vapor formed in the feed tank is merely recycled back to the source tank and both tanks equalize while liquid transfer freely takes place - without any instrumentation or hazard.

3. You say you can't "float" the tank; I think you mean "flood the feed tank". I maintain that you certainly can create a serious process hazard by flooding the feed tank with excess liquid Propylene. All that has to take place is that your present (unidentified) method of evacuating the flash vapors fails open and the level detection device fails to act. Both instruments are independent and this does not constitute a "double jeopardy", so this is very possible. Your PSV might be sized to handle this scenario (although I would doubt it until I saw the calculations), but the resulting liquid discharge is certainly a hazard that I know you would not tolerate nor want to happen. That's why I mention this.

4. I wouldn't use gate valves because they are simply not made for this application where you require bubble-tight closure on a very hazardous and flammable compressed gas liquid. You should use a soft-seated valve that can promise complete and sealed closure. A ball valve is the instrument of choice here. But don't forget the points about the ball valve that I mentioned. You are handling a saturated liquid with a high vapor pressure, so you must take care when you trap such a fluid - even in the bore space of a ball or plug valve.

I would point out that what you are observing are not phenomena. This is all a result of hard, practical and predictable Thermodynamics. Science is not a phenomena; it is proven, repeatable, and subject to known and definable Laws. Engineers are supposed to dominate the Thermo in order to ensure that the correct application of these known Laws will take place safely and efficiently.

I have been advised by Chris Haslego, the owner of this Forum, that all previous replies and questions on the Old Forum are available to you still. You can copy them and paste them into this New Forum for a continuing thread.

I hope this rather long explanation helps you understand how your feed controls are working and why.

Art Montemayor

[afdmello]

Thanks Art,

I will take some time to understand and I am sure ur reply will defenitely address all my doubts

Greatly indebted for ur time and effort in making me understand.

AFD

[naihusna]

I have a question regarding the propylene. Is it possible to transfer the propylene from pressurized tank (17-20 barg) into the vaporiser without pump? For your information, the vaporiser should be in 30-32 barg to maintain the propylene in liquid before it can be vaporised to the compressor. There are 2 vaporisers which will be loaded with propylene one at a time. Let's name it as A and B. A is depressurised from 30 bar g to 15 bar g to allow for propylene loading from tanker which is at 17 barg. While loading to the A, the vapour return line from B which has been pressurised to 30 bar g will be cracked open to pressurize the tanker at 21 barg. There is a pressure control valve at the vapour return line which will be set at 21 bar g. Once the level reach high in the vaporiser, the control valve at the inlet to the vaporiser will be closed. So, I would appreciate the your comment on this conceptual whether it can be implemented or not.

[Art Montemayor]

naihusna:

Your post is very difficult to understand. It isn’t the grammar; it’s the manner and logic in which you describe what you are trying to do. I frankly can’t understand what it is that you are trying to do. I think I have an idea, but I’m not sure.

Can you draw us a sketch on Excel of what it is that you are trying to configure? We need to know basically what process you are trying to employ. For example:

1. Do you have a Propylene Storage tank containing saturated liquid Propylene? If so, tell us what are the storage conditions of this tank. Is this storage tank at grade (ground level)?
2. Are you vaporizing the saturated liquid propylene in the storage tank in order to compress the resultant vapor propylene to a higher pressure to use downstream in another operation? Is this downstream operation a reaction with other chemicals?
3. What are the conditions at which you need the final, pressurized propylene gas – the temperature and pressure?

If you have propylene stored at 17 barg (18 bara), then your saturated storage temperature is 41.3 ^oC (106.3 ^oF). This would be about average in a hot environment, like the MidEast.

Normally if you have need for gaseous propylene at 30 barg (31 bara) and you have only the saturated 17 barg liquid available, then you use a pump to pump the liquid into a vaporizer that contains the liquid at the desired downstream pressure (30 barg?) and you vaporize this saturated liquid at that pressure. You continue to add 30 barg liquid on liquid level demand as you vaporize the amount that you need.

Is this something similar to what you are trying to do? If so, why not just pump the liquid into a vaporizer working at the pressure you desire (or higher)? Without an accurate, detailed description of what it is that you are trying to do, we are only guessing and converting what is probably a simple response into a very long, and drawn-out thread without specific answers. Await your response.

[naihusna]

I'm so sorry because my explanation not very clear. Generallly, what you have explained in your second last para. is the normal condition of the system. We have pump to pump the liquid propylene into the vaporiser at 30 barg. Now, we are considering the temporary propylene system to transfer liquid propylene from mobile source (i.e. iso-tanker lorry which carries propylene liquid). The iso-tanker will be parked nearby the vaporiser on the ground level whereas vaporiser is about 3 m from the ground level. This time we will not use pump to transfer the liquid from iso-tanker into the vaporiser. Only to move it by differential pressure. So, the question is, "Is it possible to transfer it without pump?" As I told you before, the operating pressure in the iso-tank is 17-20 barg with 28 deg C and the desired pressure in the vaporiser is 30-32 barg with 77 deg C (i.e. the propylene is in the liq phase). I attached also the schematic diagram for your clear picture of the system. Appreciate your comment. Sorry for my weakness of explanation.

Attached Files

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